Process for producing diphenyl carbonate

ABSTRACT

Processes for producing diaryl carbonates are disclosed, where such processes may provide for the production of diaryl carbonates from green house gases, such as carbon dioxide. The processes disclosed advantageously integrate diethyl carbonate and diaryl carbonate production, eliminating the need for solvent-based extractive distillation, as is commonly required when producing diaryl carbonates from dimethyl carbonate, providing for the integration of separation equipment and raw material usage, and reducing the operating and capital requirements for such processes. In some embodiments, processes disclosed herein may be operated essentially closed-loop with respect to ethanol usage, for example.

CROSS-REFERENCE TO RELATED APPLICATION

The present application is a Continuation-In-Part application of U.S.patent application Ser. No. 12/029,283, entitled “Process for ContinuousProduction of Organic Carbonates or Organic Carbamates and SolidCatalysts Therefore,” filed by J. Yong Ryu on Feb. 11, 2008, thecontents of which are hereby incorporated by reference in their entiretyherein.

FIELD OF THE DISCLOSURE

Embodiments disclosed herein relate generally to processes and solidcatalysts for reactions involving alcoholysis, transesterification, anddisproportionation. More specifically, embodiments disclosed hereinrelate to processes for the continuous production of organic carbonates,organic carbamates and other products via alcoholysis,transesterification, and/or disproportionation over a solid catalyst. Inparticular, embodiments disclosed herein relate to processes for theproduction of diaryl carbonates.

BACKGROUND

Transesterification, or the exchange reaction of esters with alcohols(an alcoholysis reaction), is an important class of reactions which maybe catalyzed by both acid and base catalysts. Examples oftransesterification, in general, include chemical reactions involvingorganic carbonates and carboxylic acid esters as reactants, products, orboth. Other transesterification reactions include the production ofbiodiesel by transesterification of triglycerides with ethanol ormethanol. Alcoholysis, in general, is a reaction where one or morefunctional groups of a compound are replaced by alkoxy or aryloxy groupof an alcohol (alkyl or aryl hydroxyl compound). Examples of alcoholysisinclude chemical reactions involving urea, where amine groups arereplaced by alkoxy groups to produce organic carbamates and carbonates.

Carboxylic acid esters are produced by transesterification of acarboxylic acid ester with an alcohol in the presence of acid and basecatalyst. Sulfuric acid (homogeneous) and acid resins (solid) arepreferred acid catalysts. Soluble bases, such as NaOH and KOH, variousNa/K alkoxides or amines (homogeneous), and various basic resins (solid)are preferred base catalysts. Although catalysts can be eitherhomogeneous catalyst or heterogeneous catalyst for thetransesterification of carboxylic esters, base catalysts are, ingeneral, more effective than acid catalysts. For example, long chainalkyl methacrylic esters are produced by exchange reaction of methylmethacrylate with a long chain alcohol in the presence of a basecatalyst.

Biodiesel may be produced by transesterification of vegetable oils(triglycerides) with methanol or ethanol by using a homogeneous basecatalyst, such as sodium methoxide or calcium acetate, as disclosed inU.S. Pat. Nos. 6,712,867 and 5,525,126, and a basic solid catalyst, suchas a mixed oxide of zinc oxide and alumina or zinc aluminate (zinc oxidesupported on alumina and calcined at a high temperature). Solid zincaluminate catalysts are disclosed in U.S. Pat. No. 5,908,946 and U.S.Patent Application Publication No. 2004/0034244, for example.

U.S. Pat. No. 5,908,946 discloses a two-step process for producingesters by reacting vegetable oils or animal oils with an alcohol in thepresence of solid catalysts such as zinc oxide or spinel type zincaluminates. In the first step, the conversion of triglyceride is forcedto a high conversion, usually higher than 90%. In the second step, theremaining triglycerides, diglycerides and monoglycerides are converted.The transesterifications are performed at a temperature from 230 to 245°C. at about 5.2 bar (about 725 psia). High conversion requiresrelatively low flow rates of a feed mixture (0.5 h⁻¹ or lower spacevelocity).

U.S. Pat. No. 6,147,196 discloses a process for producing high purityfatty acid esters from plant or animal oil in the presence of aheterogeneous catalyst (zinc aluminate). U.S. Patent ApplicationPublication No. 2004/0034244 relates to a processing scheme forproducing alkyl esters from vegetable or animal oil and an alcohol inthe presence of a heterogeneous catalyst (zinc aluminate). The estersare produced by transesterification in two fixed bed reactors. Highconversion of triglyceride was obtained in the first reactor. Afterseparating glycerol from the first transesterification reaction stream,the remaining unconverted triglyceride, diglyceride and monoglycerideare converted to esters in the second reactor. The transesterificationis performed at 200° C., about 62 bar (900 psia) and 0.5 h⁻¹ spacevelocity.

W. Xie et al. (J. Mol. Cat. A: Chem. 246, 2006, pp 24-32) discussmethanolysis of soybean oil in the presence of a calcined Mg—Alhydrotalcite catalyst. The calcined hydrotalcites with an Mg/Al ratio of3.0 derived from calcinations at 500° C. is a catalyst that can givehigh basicity and excellent catalytic activity for this reaction. Theyreport the soluble basicity of the hydrotalcites calcined at varioustemperatures.

Diesel engines emit more particulates and NO_(x) than gasoline engines.It is reported that dialkyl carbonates are effective in reducingparticulates in diesel engine exhaust. According to U.S. Pat. No.5,954,280, urea and ammonia are effective NO_(x) reducing agents. Butusing urea and ammonia for diesel engine has practical problems orinconveniences. U.S. Pat. No. 6,017,368 discloses ethyl carbamate aseffective at reducing NO_(x) from diesel engines. U.S. Pat. No.4,731,231 (1988) reports that sublimed cyanuric acid can be an effectiveagent for the elimination or reduction of NO_(x). High temperaturesublimation of cyanuric acid produces isocyanic acid (HNCO), which isbelieved to be responsible for elimination of NO_(x). EP 0363681 and EP0636681 disclose a carbonate ester of an aliphatic triol or tetraol as acomponent of low smoke lubricating agents.

N-aryl methyl carbamate is produced by reaction of an aromatic aminewith a dimethyl carbonate, typically in the presence of a base catalystdue to low reaction rates in the absence of a catalyst. N-aryl methylcarbamate can be decomposed to produce aromatic isocyanate at elevatedtemperature. For example, toluene dicarbamate is produced by reactingtoluene diamine with dimethyl carbonate in the presence of a catalyst.Decomposition of toluene dicarbamate at elevated temperature yieldstoluene diisocyanate.

Organic carbonates (diesters of carbonic acid) are useful compounds thatmay be used as solvents, alkylating agents, carbonylation agents,co-polymerization agents, fuel additives, etc. Dimethyl carbonate (DMC)is an important dialkyl carbonate, commonly used as a raw material forthe production of diphenyl carbonate (DPC, a diaryl carbonate). Thereare various processes for commercial production of DMC. In one suchcommercial process, DMC is produced by transesterification of a cycliccarbonate with methanol in the presence of a homogeneous catalyst.Although patents may disclose use of homogeneous catalysts orheterogeneous catalysts for transesterification of a cyclic carbonatewith methanol, there is currently no commercial practice where aheterogeneous or solid catalyst is used for the production of DMC,likely due to the short cycle length of heterogeneous catalysts for suchprocesses. DPC is commonly co-polymerized with a diol, such as bisphenolA, to produce polycarbonates. Polycarbonates are used in various specialapplications such as memory disks, windshields, engineering plastics,optical materials, etc.

Current techniques for production of diaryl carbonates using anon-phosgene process produces aromatic carbonates, such as DPC, bytransesterification of DMC with phenol to produce methyl phenylcarbonate and methanol, followed by disproportionation of the methylphenyl carbonate to produce DPC and DMC in the presence of homogeneousorganometallic catalysts by employing a series of multiple reactivedistillation reactors. A preferred homogeneous catalyst is titaniumalkoxide. Such processes are disclosed in U.S. Pat. Nos. 4,045,464,4,554,110, 5,210,268, and 6,093,842, for example. The homogeneouscatalysts are recovered from the heaviest portion of the product streamsas a solid, which may then be converted to soluble homogeneous catalystto recycle.

Use of a homogeneous catalyst in the production of DPC often requiresseparation of the homogeneous catalyst from the product, especiallywhere the catalysts are used at relatively high feed rates. To alleviatethis and other shortcomings associated with using homogeneous catalystsfor the production of diaryl carbonates, U.S. Pat. Nos. 5,354,923 and5,565,605, and PCT Application Publication WO03/066569 disclosealternative processes where heterogeneous catalysts are used. Forexample, U.S. Pat. No. 5,354,923 discloses titanium oxide catalysts inpowder form to demonstrate the preparation of EPC, MPC and DPC from DECor DMC and phenol. U.S. Pat. No. 5,565,605 discloses microporousmaterials containing Group 4 elements as the catalysts fortransesterification and disproportionation. However, solid catalysts inpowder form are typically unsuitable or less preferable for large volumecommercial production of DPC or methyl phenyl carbonate. WO03/066569discloses a process for continuously producing DPC in the presence of aheterogeneous catalyst prepared by supporting titanium oxide on silicain a two-step fixed bed process by reacting DMC with phenol.

Z-H Fu and Y. Ono (J. Mol. Catal. A. Chemical, 118 (1997), pp. 293-299)and JP Application No. HEI 07-6682 disclose heterogeneous catalysts forthe preparation of diphenyl carbonate by transesterification of DMC withphenol to MPC and disproportionation of MPC to DPC in the presence ofMoO₃ or V₂O₅ supported on an inorganic support such as silica, zirconia,or titania. The transesterification and disproportion are carried out ina reactor-distillation tower consisting of a reactor and distillationtower with removal of the co-products by distillation.

U.S. Patent Application Publication Nos. 2007/0093672 ('672) (now U.S.Pat. No. 7,378,540) and 2007/0112214 ('214) (now U.S. Pat. No.7,288,668) disclose processes for producing various organic carbonates,such as diaryl carbonates, including DPC, in the presence ofheterogeneous catalysts. In the '214 publication, the necessaryreactions (transesterification and disproportionation) are performed inliquid phase in the presence of a heterogeneous catalyst. Multiple fixedbed reactors for the transesterification and disproportionationreactions are connected to a single distillation column, where lightcompounds such as ethanol and DEC are removed as an overheads fraction,and the higher boiling compounds, including DPC, are removed as a mixedbottoms fraction. DPC is then recovered from the bottoms fraction.

The '672 publication discloses a process for making diaryl carbonatesand dialkyl carbonates by performing the necessary reactions in adual-phase (vapor and liquid) reaction over various solid catalysts fortransesterification and disproportionation. The chemical reactionsproducing organic carbonates are performed in a series of fixed bedreactors, while performing separation of light co-product in liquidphase to vapor phase in order to shift the unfavorable equilibriumreaction toward the desired product. The process is especially usefulfor the production of alkyl aryl carbonates such as EPC (ethyl phenylcarbonate) and diaryl carbonates such as DPC (diphenyl carbonate). Theprocess is also useful for the production of dialkyl carbonates such asDEC. A series of fixed bed reactors are connected at different positionson a single distillation column via side-draw streams and returnstreams. The distillation column also contains separation stages abovethe last reactor in the series and below the first reactor in theseries. The heterogeneous catalysts may be prepared by depositing one ortwo metal oxides of Ti, Zr, Nb, Hf, Ta, Mo, V, Sb, etc. on poroussupports, such as silica gel. The heterogeneous catalysts may also beprepared by grafting one or more organometallic compounds from theelements of Ti, Zr, Nb, Hf, Ta, Mo, V, Sb, etc. on a porous support,which has surface hydroxyl groups or a mixture of hydroxyl and alkoxygroups.

Various other processes for the production of organic carbonates withheterogeneous catalysts are disclosed in U.S. Pat. Nos. 5,231,212,5,498,743, and 6,930,195.

P. Ball et al. (C₁ Mol. Chem. Vol. 1, 1984, pp. 95-108) studied thechemistry of dialkyl carbonate production in the presence of varioushomogeneous or heterogeneous catalysts. For example, dimethyl carbonateis produced by alcoholysis of urea. Dibutyltin dimethoxide is reportedas a particularly effective catalyst. It is reported that heterogeneouscatalysts are also effective for the chemistry in the presence ofco-catalysts, such as 4-dimethylaminopyridine and PPh₃. The reportedheterogeneous catalysts are Al₂O₃, Sb₂O₃, and silica. Fused SiO₂ is nota catalyst, but becomes catalytic in the presence of PPh₃.

In U.S. Pat. No. 7,074,951, dialkyl carbonate is produced by alcoholysisof urea with an alcohol in the presence of a homogeneous tin complexcatalyst in the presence of a high boiling electron donor atomcontaining solvent, such as triglyme. This patent also demonstrates thecapability of producing DMC continuously for about 1500 hours.

EP 1629888 and D. Wang et al. (Fuel Processing Tech. 88, 8, 2007, pp807-812) disclose that DMC and DEC may be produced in the presence ofzinc oxide and zinc oxide supported on silica. These publications arecompletely silent about the catalyst stability or catalyst cycle length.

Catalyst deactivation during transesterification and disproportionationreactions may be caused by the deposition of heavy polymers on thecatalyst surface and pores. The catalyst deactivation rate by polymerdeposition increases with the concentration of alkyl aryl carbonate anddiaryl carbonate or both in a reaction mixture. Depolymerization ofpolymers on the heterogeneous catalysts is disclosed in the '672publication. However, depolymerization may results in only a partialrecovery of solid catalyst activity.

U.S. Pat. Nos. 6,768,020 and 6,835,858 disclose processes for makingdialkyl carbonates and co-product propylene glycol by reaction ofpropylene carbonate with DMC, water, or both, in the presence of solidcatalyst such as lanthanum oxide and zinc oxide supported on alumina,silica, etc. Catalyst instability is partially solved in U.S. Pat. No.6,768,020 by depositing a large amount of lanthanum oxide on a supportsuch as alumina and silica.

A favored technique to compensate for catalyst deactivation is theramping up of the reaction temperature as the catalyst deactivates. Thistechnique, unfortunately, often accelerates deactivation ofheterogeneous catalysts.

Long, stable performance of a solid catalyst is generally required forcommercial production using a heterogeneous catalyst. Catalyst costs,downtime associated with catalyst replacement, and other factors asknown in the art dictate that heterogeneous catalysts have a minimumlifespan, typically greater than 3 months, 6 months, or a year,depending upon the process.

Although heterogeneous catalysis of various transesterificationreactions is possible, as described by the various patents andpublications above, they do not report longevity or cycle length of thecatalyst. It is the experience of the present inventor that suchheterogeneous catalysts have undesirably short cycle lengths.

Accordingly, there exists a need for transesterification and/ordisproportionation processes using heterogeneous catalysts with improvedcatalyst performance.

SUMMARY OF THE DISCLOSURE

In one aspect, embodiments disclosed herein relate to an alcoholysisprocess, the process including: feeding reactants and a trace amount ofsoluble organometallic compound to a reactor comprising a solidalcoholysis catalyst; wherein the soluble organometallic compound andthe solid alcoholysis catalyst each independently comprise a Group II toGroup VI element. The solid catalyst and the organometallic compound mayinclude the same Group II to Group VI element in some embodiments.

In another aspect, embodiments disclosed herein relate to a process forthe production of dialkyl carbonates, the process including: feeding analcohol, and an alcoholysis reactant comprising at least one of urea, anorganic carbamate, and a cyclic carbonate to a first reaction zonecomprising a solid alcoholysis catalyst; feeding a solubleorganometallic compound to the first reaction zone, wherein the solidalcoholysis catalyst and the soluble organometallic compound eachindependently comprise a Group II to Group VI element.

In another aspect, embodiments disclosed herein relate to a process forthe production of diaryl carbonate, the process including: feeding anaromatic hydroxy compound and a dialkyl carbonate to a first reactionzone comprising a solid transesterification catalyst; and feeding asoluble organometallic compound to the first reaction zone, wherein thesolid transesterification catalyst and the soluble organometalliccompound each independently comprise a Group II to Group VI element.

In another aspect, embodiments disclosed herein relate to a process forthe production of an alkyl aryl carbonate, the process including:feeding an aromatic hydroxy compound and a dialkyl carbonate to a firstreaction zone comprising a solid transesterification catalyst; andfeeding a soluble organometallic compound to the first reaction zone,wherein the solid transesterification catalyst and the solubleorganometallic compound each independently comprise a Group II to GroupVI element.

In another aspect, embodiments disclosed herein relate to a process forthe production of biodiesel, the process including: feeding an alcoholand a glyceride to a first reaction zone comprising a solidtransesterification catalyst; and feeding a soluble organometalliccompound to the first reaction zone, wherein the solidtransesterification catalyst and the soluble organometallic compoundeach independently comprise a Group II to Group VI element.

In another aspect, embodiments disclosed herein relate to a process forthe production of an alkyl aryl carbonate, the process including:feeding an aromatic hydroxy compound and a dialkyl carbonate to a firstreaction zone comprising a solid transesterification catalyst; andfeeding a soluble organometallic compound to the first reaction zone,wherein the solid transesterification catalyst and the solubleorganometallic compound each independently comprise a Group II to GroupVI element.

In another aspect, embodiments disclosed herein relate to a process forthe production of biodiesel, the process including: feeding an alcoholand a glyceride to a first reaction zone comprising a solidtransesterification catalyst; and feeding a soluble organometalliccompound to the first reaction zone, wherein the solidtransesterification catalyst and the soluble organometallic compoundeach independently comprise a Group II to Group VI element

In another aspect, embodiments disclosed herein relate to a process forreactivating a spent solid alcoholysis catalyst, the process including:removing polymeric materials deposited on the catalyst; andre-depositing catalytically active metals on the solid catalyst.

In another aspect, embodiments disclosed herein relate to a process forproduction of diaryl carbonate, including: reacting an epoxide andcarbon dioxide in a first reaction zone to form first reaction productcomprising a cyclic carbonate; transesterifying the cyclic carbonatewith ethanol in the presence of a first transesterification catalyst ina second reaction zone to form a second reaction product comprisingdiethyl carbonate and glycol; separating the second reaction product torecover a first diethyl carbonate fraction and a first glycol fraction;transesterifying at least a portion of the first diethyl carbonatefraction with an aryl hydroxy compound in the presence of a secondtransesterification catalyst in a third reaction zone to form a thirdreaction product comprising ethyl aryl carbonate and ethanol; separatingthe third reaction product to recover an ethyl aryl carbonate fractionand a first ethanol fraction; disproportionating at least a portion ofthe ethyl aryl carbonate fraction in the presence of adisproportionation catalyst in a fourth reaction zone to form a fourthreaction product comprising diaryl carbonate and diethyl carbonate;separating the fourth reaction product to recover a diaryl carbonatefraction and a second diethyl carbonate fraction; recycling at least aportion of the first ethanol fraction to the second reaction zone; andrecycling at least a portion of the second diethyl carbonate fraction tothe third reaction zone.

In another aspect, embodiments disclosed herein relate to a process forproducing diaryl carbonate, including: reacting ammonia and carbondioxide in a first reaction zone to form first reaction productcomprising a urea; transesterifying the urea with ethanol in thepresence of a first transesterification catalyst in a second reactionzone to form a second reaction product comprising diethyl carbonate andammonia; separating the second reaction product to recover a firstdiethyl carbonate fraction and a first ammonia fraction;transesterifying at least a portion of the first diethyl carbonatefraction with an aryl hydroxy compound in the presence of a secondtransesterification catalyst in a third reaction zone to form a thirdreaction product comprising ethyl aryl carbonate and ethanol; separatingthe third reaction product to recover an ethyl aryl carbonate fractionand an ethanol fraction; disproportionating at least a portion of theethyl aryl carbonate fraction in the presence of a disproportionationcatalyst in a fourth reaction zone to form a fourth reaction productcomprising diaryl carbonate and diethyl carbonate; separating the fourthreaction product to recover a diaryl carbonate fraction and a seconddiethyl carbonate fraction; recycling at least a portion of the ethanolfraction to the second reaction zone; and recycling at least a portionof the second diethyl carbonate fraction to the third reaction zone.

Other aspects and advantages will be apparent from the followingdescription and the appended claims.

BRIEF DESCRIPTION OF DRAWINGS

FIG. 1 is a simplified process flow diagram illustrating a process forthe production of diaryl carbonates according to embodiments disclosedherein.

FIG. 2 is a simplified process flow diagram illustrating a process forthe production of diaryl carbonates according to embodiments disclosedherein.

FIG. 3 is a simplified process flow diagram illustrating a process forthe production of diaryl carbonates according to embodiments disclosedherein.

FIG. 4 is a graphical representation of transesterification using ahomogeneous catalyst.

FIG. 5 is a graphical representation of catalyst activity followingcatalyst reactivation according to embodiments disclosed herein.

FIG. 6 is a graphical representation of solid catalyst activity when atrace amount of soluble organometallic compound is added to the reactoraccording to embodiments disclosed herein.

FIG. 7 graphically compares heterogeneous catalyst activity with solidcatalyst activity when a trace amount of soluble organometallic compoundis added to the reactor according to embodiments disclosed herein.

FIG. 8 is a graphical representation of solid catalyst activity when atrace amount of soluble organometallic compound is added to the reactoraccording to embodiments disclosed herein.

FIGS. 9A and 9B are a graphical representation of solid catalystactivity during production of EPC and DPC, respectively, when a traceamount of soluble organometallic compound is added to the reactoraccording to embodiments disclosed herein.

FIG. 10 is a graphical representation of heterogeneous catalyst activityduring production of DPC where the catalyst is grafted simultaneouslywhile performing a transesterification reaction.

FIG. 11 graphically illustrates the conversion of EPC to DPC and DEC inthe absence of solid catalysts according to embodiments disclosedherein.

FIG. 12 graphically presents results from the alcoholysis of propylenecarbonate with ethanol to produce DEC and propylene glycol in thepresence of a solid catalyst according to embodiments disclosed herein.

FIG. 13 presents results from the production of DEC using a homogeneouscatalyst.

FIG. 14 presents results from the production of DEC using a solidcatalyst according to embodiments disclosed herein.

FIG. 15 is a simplified process flow diagram for the production ofdialkyl carbonates using a solid catalyst according to embodimentsdisclosed herein.

FIG. 16 presents results from the production of DEC from ethyl carbamateusing a solid catalyst according to embodiments disclosed herein.

FIG. 17 presents results from the alcoholysis of canola oil withmethanol using a solid catalyst according to embodiments disclosedherein.

FIG. 18 is a simplified process flow diagram for the continuousproduction of DEC and propylene glycol co-product by performingalcoholysis of propylene carbonate with ethanol in the presence of asolid catalyst according to embodiments disclosed herein.

FIG. 19 is a simplified block flow diagram illustrating the process forthe production of diphenyl carbonate (DPC) according to embodimentsdisclosed herein.

FIG. 20 is a simplified block flow diagram illustrating the integratedprocess for the production of diphenyl carbonate (DPC) according toembodiments disclosed herein.

FIG. 21 graphically presents results from the alcoholysis of propylenecarbonate with ethanol to produce DEC and propylene glycol in thepresence of a solid catalyst according to embodiments disclosed herein.

FIG. 22 presents results from the production of DEC from ethyl carbamateusing a solid catalyst according to embodiments disclosed herein.

FIG. 23 graphically illustrates results from the catalytictransesterification of DEC with phenol to produce EPC, an intermediateproduct in the DPC production process.

FIG. 24 graphically illustrates results from the catalyticdisproportionation of EPC to produce DPC and DEC.

DETAILED DESCRIPTION

In one aspect, embodiments disclosed herein relate to alcoholysis,transesterification, and/or disproportionation processes using solidcatalysts. As used herein, alcoholysis is termed to represent variouschemical reactions where an organic hydroxyl compound (alcohol) isinvolved as one of two reactants to produce a product and a co-product.Alcoholysis may be defined as breaking bonds (C—Y) between a carbon atomand a heteroatom Y of molecules by an alcohol molecule (ROH).Alcoholyses are reactions involving carbonyl groups of a molecule andthe carbonyl group itself is retained in the product molecule.Therefore, the C atom of the C—Y bond is the carbon atom of carbonylgroup of a molecule. Generally alcoholysis is a reversible reaction, andmay be represented as follows:

where Y is a heteroatom or a heteroatom of a functional group, and R^(b)is alkyl, aryl, or a functional group having one or more heteroatoms.

Examples of alcoholysis reactions are the reaction of an alcohol withdiesters of carbonic acid, esters of carboxylic acids, urea, andcarbamates. Alcoholysis of a dialkyl carbonate (often referred to astransesterification in literature) with phenol produces alkyl arylcarbonate and an alcohol. Alcoholysis of an ester of carboxylic acidwith an alcohol exchanges the alkyl group of the ester with the alkylgroup of the alcohol molecule and produces a new alcohol molecule.Alcoholysis of urea with an alcohol produces an organic carbamate andammonia. Alcoholysis of an organic carbamate with an alcohol produces adialkyl carbonate and ammonia. Specific examples of alcoholysisreactions are the transesterification of DEC with phenol to produce EPCand ethanol, alcoholysis of urea or organic carbamate with an alcohol toproduce organic carbamate or dialkyl carbonate and ammonia,transesterification of triglyceride with methanol to produce methylesters (biodiesel) and glycerin.

Although disproportionation of unsymmetrical carbonic acid diesters andthe reaction of a dialkyl carbonate with an organic amine do not involvean alcohol as reactant, it is considered here that these type ofreactions are also defined as alcoholysis reactions as well forconvenience, because RA groups (R is alkyl or aryl, and A is oxygen atomor nitrogen atom) are involved in the reaction mechanisms at molecularlevel. Therefore, transesterification and disproportionation are used assynonyms to alcoholysis as necessary for the description of variousembodiments. A few of the alcoholysis reactions mentioned above may berepresented by the following reactions:

In another aspect, embodiments disclosed herein relate to a noveltechnique for maintaining catalyst activity for a solid catalyst for anextended cycle time. The cycle time or cycle length of a solid catalystherein is defined to be the time period over which a solid catalyst cancontinuously be used without interruption for an intended chemicalreaction. For example, if a catalyst requires catalyst regeneration orreplacement after continuous use of 6 months, the catalyst cycle lengthor time is 6 months. According to technique disclosed herein, solidcatalysts for alcoholysis processes may retain catalyst activity for anextended cycle time, such as greater than 3 months, 6 months, 1 year,1.5 years, and 2 years or more, in various embodiments.

During the transesterification of DEC with phenol, deactivation ofheterogeneous catalysts (a titanium oxide and a mixed oxide of niobiumand titanium oxide supported on silica) was observed by the presentinventor and reported in Test 4 of the '672 publication.Depolymerization of polymer buildup on the catalyst to improve catalystactivity was also demonstrated in Test 6B of the '672 publication.However, the catalyst regeneration by depolymerization resulted in onlya partial recovery of the original catalyst activity. The nature ofcatalyst deactivation was not fully understood at that time.

It has been surprisingly found that heterogeneous transesterificationcatalysts, such as heterogeneous catalysts for making DPC, deactivatedue to two primary causes: polymer deposition and leaching ofcatalytically active metal component. Heterogeneous catalysts for makingdialkyl carbonate by transesterification of a cyclic carbonate with analcohol deactivate primarily due to leaching of catalytically activemetal components.

During alcoholysis or transesterification reactions over heterogeneouscatalysts, the catalytically active metal components on the solidcatalysts may leach out of the heterogeneous metal oxide catalysts andorganometallic catalysts immobilized on various porous supports into thereaction medium under reaction conditions, resulting in permanentcatalyst deactivation. This results in an unacceptably short catalystlife for commercial heterogeneous catalysts that may be used for thecontinuous production of various organic carbonates. Additionally, asmentioned above, polymer deposition may also affect the performance of atransesterification catalyst. A further mode of catalyst deactivation ispoisoning.

A heterogeneous catalyst to be used in a commercial fixed bed reactormust have reasonable longevity for both cycle time and total servicetime. Absent poisoning, and if there is no or little deposition ofpolymers on a heterogeneous catalyst, the rate of dissolution of anactive metal component from a heterogeneous catalyst may determine thelongevity of the catalyst.

Embodiments disclosed herein relate to processes for maintaining aconstant or near constant solid catalyst activity for an extended periodof time acceptable for continuous production of various organiccompounds at a commercial scale. Such processes may be particularlyuseful for the continuous production of various organic carbonates, suchas diaryl carbonates, dialkyl carbonates, and alkyl aryl carbonates, aswell as in other transesterification reactions, such as for theproduction of biodiesel. Selected embodiments disclosed herein relatesto processes for maintaining stable catalyst activity for extendedperiod time for large commercial reactors for continuous production oforganic carbonates, carboxylic acid esters, or organic carbamates.

The novel technique for maintaining catalyst activity for a solidcatalyst for an extended cycle time is that adding a trace amount ofsoluble active metal components to the liquid feed stream to a solidcatalyst-containing reactor may result in a constant or near constantcatalyst activity over extended periods of cycle time. It wasunexpectedly found that adding a trace amount of soluble active metalcomponents to the liquid feed stream to a solid catalyst-containingreactor may effectively counter balance the metal loss due to the metalleaching out of the solid catalysts, such as by redepositing activemetal component on the catalyst, resulting in a constant or nearconstant catalyst activity over extended periods of cycle time. Forexample, current efforts by the present inventor indicate that catalystactivity may be maintained for greater than one year by adding a traceamount of soluble active metal compounds to liquid feed stream fed tofixed bed or moving bed reactors containing a solid catalyst.

The amount of active metal compound required to maintain activity of asolid catalyst may range from less than 1 ppm to about 3000 ppm,depending upon the specific active metal component, the reactants, andother feed components. The amount of active metal in the feed, such asfor the production of organic carbonates by transesterification, may beone, two, or more orders of magnitude lower than the concentration ofhomogeneous catalyst fed for a comparable process using homogeneouscatalyst only. In some embodiments, the active metal compound may be fedat a rate of 1 to 400 ppm by weight; 10 to 300 ppm by weight in otherembodiments; 15 to 200 ppm by weight in other embodiments; 20 to 150 ppmby weight in other embodiments; and from 30 to 100 ppm by weight in yetother embodiments, based on the total weight of the liquid entering thecatalytic reaction zone.

For example, where a solid catalyst includes an active metal, such as aGroup II-VI metal, a trace amount of a soluble organometallic compoundhaving the same Group II-VI active metal may be fed to the reactor tomaintain the activity of the solid catalyst. As a specific example,where a solid catalyst includes titanium as an active metal, a solubleorganometallic compound having titanium may be used.

Where a series of reactors is used, such as a transesterificationreactor in series with a disproportionation reactor, a trace amount ofsoluble organometallic compound may be fed to one or both reactors tomaintain catalyst activity in the respective reactors. In someembodiments, by feeding a trace amount of soluble organometalliccompound to only the first reactor in a series of reactors, solidcatalyst activity may be maintained in each of the reactors. Forexample, where a transesterification reactor includes a solid mixedoxide catalyst of titanium and niobium, and a disproportionation reactorcontains a solid titanium alkoxide grafted on silica, adding a traceamount of soluble organotitanium compound or a soluble titaniumcompound, such as titanium oxyalkoxide, to the first reactor, the cycletime of both types of solid catalysts may be extended.

The soluble organometallic compound may be recovered and recycled, ifdesired. In some embodiments, it may not be economical to recover theactive metal from the reactor effluent streams for recycle. Whenrecovered, the active metal component in the reactor effluent stream maybe recovered from the heavy bottom stream as a solid material andconverted to a soluble organometallic compound, which may be recycled tothe reactor, such as by reacting the recovered solid material withorganic carbonate or a mixture of organic carbonate and alcohol at anelevated temperature. The recovered organometallic compound, forexample, may be a metal alkoxide, metal alkoxy alkyl carbonate (metalsalt of carbonic acid monoester), metal oxyalkoxide, or a mixturethereof.

The extended alcoholysis solid catalyst life thus attained may result incommercially viable solid catalyst processes for the production oforganic carbonates, among alcoholysis and/or other transesterificationprocesses. Significant savings may be realized due to extended catalystscycle time and decreased separation requirements (fewer unit operations,resulting in potential capital and operating cost savings).

Polymer deposition on solid catalysts may also cause loss of catalystactivity. In such a case, a deactivated catalyst may be regenerated bydepolymerization techniques disclosed herein and in U.S. PatentApplication Publication No. 2007/0093672. Depolymerizatoin also cancause metal loss. In the case of following depolymerization whereheterogeneous catalyst does not recover the catalyst activity to anacceptable level of original activity, the heterogeneous catalyst mayrequire metal redeposition, such as by catalyst reactivation techniquesdisclosed herein.

Whenever there is catalyst deactivation caused by both polymerdeposition and metal leach, the catalyst activity may be restored by thecatalyst regeneration and reactivation techniques disclosed herein. Thecatalyst reactivation consists of two steps: depolymerization andsurface conditioning in the first step and metal redeposition in thesecond step. In the first step, the deactivated solid catalyst issubjected to depolymerization to remove polymers on the solid catalystand then surface conditioning by drying. In the second step,redeposition of active metal component is performed to compensate formetal loss. The reactivation of deactivated catalyst will be discussedin more detail later.

Where catalyst reactivation and/or regeneration are considered, it maybe beneficial to have multiple reactors in parallel so as to allow forcontinuous production during the catalyst reactivation and restorationprocesses.

As described above, solid alcoholysis, transesterification, anddisproportionation processes disclosed herein may include feedingreactants and a trace amount of soluble active metal compounds to areactor containing a solid catalyst, and contacting the reactants in thepresence of the solid catalyst to alcoholyze, transesterify, ordisproportionate at least a portion of the reactants. Such alcoholysis,transesterification, or disproportionation processes may include, forexample, reactions for the production of dialkyl carbonates, diarylcarbonates, alkyl-aryl carbonates, biodiesels, organic esters, andN-aryl alkyl carbamates, among other reactions.

Although described with respect to alcoholysis, transesterification, anddisproportionation reactions in general above, the extension of suchprocesses to the production of organic carbonates is detailed below.U.S. Patent Application Publications 2007/0093672 ('672) and2007/0112214 ('214), as noted above, disclose processes for theproduction of organic carbonates using heterogeneous catalysts. Each ofthese is hereby incorporated by reference.

Organic Carbonate and Organic Carbamate Production

Organic carbonates or organic carbamates may be continuously produced byusing a single or multiple reactor systems in the presence of a solidcatalyst or two different solid catalysts. The solid catalyst orcatalysts require adding a trace amount of soluble active metal compoundinto the feed stream of the reactor to obtain an extended catalyst cycletime. Solid catalysts may be any physical shape, and may include variousorganometallic compounds immobilized on porous supports, and/or oxidescontaining an element or multiple elements of Group II, III, IV, V andVI supported on a suitable porous support. The catalysts may be eitheracid catalysts or base catalysts. The total amount of catalyticallyactive metal or metal components on a supported catalyst may range fromabout 0.02 wt % to about 20 wt % in some embodiments; from about 0.05 wt% to about 10 wt % in other embodiments. Another type of solid catalyticmaterials useful in embodiments disclosed herein is metal-organicframeworks (MOFs), which comprise one or more elements from Group II toVI and organic frameworks. MOFs may serve as both solid catalysts andcatalyst supports according to various embodiments.

The reactors used in embodiments disclosed herein may include anyphysical devices or a combination of two or more devices. The reactorsmay have various internal devices for vapor-liquid separation andvapor/liquid traffic.

By adding a trace amount of soluble active metal compound to a feedstream, stable catalyst activity may be maintained for surprisingly longcycle times. For example, addition of trace amounts of a soluble activemetal compound into streams fed to a fixed bed reactor to producemixtures of ethyl phenyl carbonate and diphenyl carbonate may results ina cycle time of more than 14 months on stream time. Such stable catalystperformance may result in higher productivity of a desired product. Inembodiments having a series of reactors, a trace amount of an activemetal component may be added only to the feed stream to the firstreactor. For a parallel multiple reactor system, a trace amount of anactive metal component may be added to all reactors.

The active metal components may include a compound or a mixture ofcompounds containing one or more metals of Group II, III, IV, V and VIof the Periodic Table. Examples of active metals include Mg, Ca, Zn, La,Ac, Ti, Zr, Hf, V, Nb, Ta, Cr, Nb, W, Sn, Pb, Sb, etc. The active metalcompound should be soluble in the reaction mixture or, at least, form anemulsion/colloidal solution. The amounts of trace metal in the feedstream may be sufficiently low so as to be economically unnecessary torecover metal from process stream to recycle, although one may choose todo so.

If necessary, a deactivated catalyst in a reactor may be reactivated insitu in a relatively short time, so as to be ready to replace anotherreactor in service or to restart service. Therefore, embodiments of theprocesses disclosed herein may require a spare reactor, depending on thecycle length of the catalyst and other factors.

Processes disclosed herein may be particularly useful for the continuousproduction of diaryl carbonates, such as diphenyl carbonate (DPC), alkylaryl carbonates, such as ethyl phenyl carbonate (EPC), or dialkylcarbonates, such as diethyl carbonate (DEC) or dimethyl carbonate (DMC).The reaction for producing diaryl carbonate may be performed in aplurality of reaction zones, such as a first and a second reaction zone.The first reaction zone serves to perform primarily transesterificationof a dialkyl carbonate with an aromatic alcohol to produce an alkyl arylcarbonate, although a small amount of diaryl carbonate may also beproduced. The second reaction zone serves to perform disproportionationof an alkyl aryl carbonate to produce diaryl carbonate and dialkylcarbonate. The presence of a solid catalyst in the second reaction zoneis not necessary, although one may choose to use a solid catalyst.

Dialkyl carbonates, such as DMC or DEC, may be produced by performingtransesterification of a cyclic carbonate, such as propylene carbonateor ethylene carbonate, with methanol or ethanol in a similar manner. Thereactions producing diaryl carbonate and dialkyl carbonate are performedin a multiple reactor system with material separation units to recoverproducts from reaction mixtures. Unreacted reactants and intermediatesmay be recovered for recycle or finished by performing a seconddisproportionation or a second transesterification. Unreacted phenol inthe liquid reaction mixture from transesterification zone may beseparated either prior to performing disproportionation of alkyl phenylcarbonate or after performing disproportionation. Additionally, thereare various options for purging the by-product alkyl phenyl ether fromthe reaction system. Proper arrangement of reactors with materialseparation units is within the knowledge of those of ordinary skill inthe art.

The reactions are preferably carried out as a mixed phase system, wherethe reactants and products are liquid and vapor, to shift theequilibrium to the desired direction. Alternatively, one may perform areaction in liquid phase, such as where there is no or little advantagein shifting equilibrium reaction due to higher boiling points of thereaction products than a preferred range of temperatures for performingthe reaction.

Embodiments disclosed herein may also be useful in producing organiccarbonates, such as ethyl phenyl carbonate, methyl phenyl carbonate, anddiphenyl carbonate, by performing transesterification of dialkylcarbonates, such as diethyl carbonate or dimethyl carbonate, withphenol, and disproportionation of an alkyl aryl carbonate, such as ethylphenyl carbonate or methyl phenyl carbonate, to produce diphenylcarbonate.

Embodiments disclosed herein may also be useful in producing dialkylcarbonate, such as dimethyl carbonate or diethyl carbonate, bytransesterification of a cyclic carbonate with an alcohol. In otherembodiments for the production of dialkyl carbonates, dialkyl carbonatemay be produced by alcoholysis of urea with an alcohol in the presenceof a solid catalyst. For example, in U.S. Pat. No. 7,074,951, dialkylcarbonate is produced by using a homogeneous organotin complex catalystin the presence of a high boiling electron donor atom-containingsolvent; such a process may be performed over solid catalysts accordingto embodiments disclosed herein. Various organic carbamates, such as anN-aryl alkyl carbamate, may also be advantageously produced by reactinga dialkyl carbonate with an aromatic amine in the presence of a solidcatalyst according to embodiments disclosed herein.

Any type of reactors may be used to carry out the reactions describedherein. The examples of reactors suitable for carrying out the reactionsinvolving organic carbonate or organic carbamates reactions may includedistillation column reactors, divided wall distillation column reactors,traditional tubular fixed bed reactors, bubble column reactors, slurryreactors equipped with or without a distillation column, pulsed flowreactors, catalytic distillation columns wherein slurry solid catalystsflow down the column, or any combination of these reactors.

Multiple reactor systems useful in embodiments disclosed herein mayinclude a series of multiple reactors or multiple reactors in parallelfor the first reaction zone. If a product is produced from reactants viaan intermediate product, such as an alkyl aryl carbonate, the firstreaction zone may serve to primarily produce the intermediate, althougha minor amount of final reaction product may simultaneously be producedin the first reaction zone.

The process stream from the first reaction zone, after stripping off anyalcohol and dialkyl carbonate, enters to the second reaction zone, wherediaryl carbonate is produced along with co-product dialkyl carbonate.While stripping lighter reaction product from the catalytic reactionzone, the transesterification may be simultaneously performed, shiftingthe equilibrium reaction toward the forward reaction.

The reactions producing organic carbonates or organic carbamates aretypically performed at a temperature in the range from about 104° C. toabout 260° C. (about 220° F. to about 500° F.) in some embodiments; from121° C. to about 232° C. (about 250° F. to about 450° F.) in otherembodiments. The pressure for a reaction depends on boiling points ofthe reactants and products, the type of reactor to be used, and whetherliquid or dual phase (vapor/liquid) exists in the reaction zone. Ingeneral, reactor pressures may be in the range from sub-atmosphericpressure to about 22 bar (about 319 psia) in some embodiments; and fromabout 0.005 bar to about 17 bar (0.1 psia to about 250 psia) in otherembodiments. In a class of embodiments, reactions may be performed usinga suitable solvent which does not interfere with separation of reactionproducts.

In selected embodiments, embodiments disclosed herein are particularlyuseful for the continuous production of diaryl carbonates from a dialkylcarbonate and an aromatic hydroxy compound, such as the production ofdiphenyl carbonate (DPC) from a dialkyl carbonate and phenol. One routefor producing DPC is the reaction of diethyl carbonate (DEC) with phenolin the presence of one or more solid catalysts. The advantages ofproducing DPC by using DEC may include energy saving and material savingfor the construction of a plant because separation of materials from anazeotrope is not necessary. All materials need energy to produce them.Thus, saving construction materials, and energy, is considered to be“green.” In contrast, the current commercial non-phosgene process ofproducing DPC uses DMC as one of raw materials. DMC and methanol have tobe separated from an azeotrope-forming process stream by solventextractive distillation. Operating extractive distillation units isenergy intensive. Although production of DPC via DMC is possible, use ofDEC may be preferred due to the energy and material savings.

Embodiments disclosed herein may also be useful for producing dialkylcarbonates by transesterification of a cyclic carbonate with an alcohol,such as ethanol or methanol.

Producing DPC from a dialkyl carbonate and phenol involves two reactionsteps; transesterification in a first reaction zone, followed bydisproportionation in a second reaction zone. The reactions may beillustrated as follows:

where the net reaction may be illustrated as:(C_(n)H_(2n+1)O)₂C═O+2C₆H₅OH

(C₆H₅O)₂C═O+2C_(n)H_(2n+1)OH  (3)

Reaction (1) is the transesterification of a dialkyl carbonate withphenol to produce alkyl phenyl carbonate and alcohol. Reaction (2)involves disproportionation of alkyl phenyl carbonate to producediphenyl carbonate and dialkyl carbonate. Both reaction steps areequilibrium reactions. However, the disproportionation isthermodynamically much more favorable than the transesterification. Thetransesterification is primarily performed in the first reaction zone,which may include a single reactor or a multiple reactor system. Thedisproportionation reaction may then be primarily performed in a secondreaction zone.

Producing a dialkyl carbonate by transesterification of a cycliccarbonate with an alcohol is also two-step equilibrium reaction. Bothacidic and basic catalysts may be used for transesterification of acyclic carbonate with an alcohol.

According to embodiments disclosed herein, to obtain prolonged catalystcycle length, a trace amount of a soluble metal compound is added to thereactor feed stream. For transesterification of cyclic carbonate with analcohol to produce a dialkyl carbonate and a diol, either a solid baseor acid catalyst may be used. One may also perform thetransesterification by substituting a portion of alcohol with water.Alternatively transesterification may be performed in the first step,followed by the reaction of unconverted cyclic carbonate andintermediate with a water-alcohol mixture to produce a glycol as a majorreaction product in the second step. Addition of water substantiallyincreases the conversion of cyclic carbonate or productivity of a diol.However, the water advantage is only realized in a reduced yield ofdialkyl carbonate.

Catalysts Useful for Organic Carbonate and Organic Carbamate Production

As described above, catalysts useful for organic carbonate and organiccarbamate production may include supported solid catalysts having one ormore active metals from Groups II, III, IV, V and VI of the PeriodicTable. One type of catalyst useful in embodiments disclosed hereinincludes an organometallic compound or multiple organometallic compoundsof the above elements immobilized on a porous support. Porous supportsuseful in embodiments disclosed herein may include surface functionalgroups such as hydroxyl groups, alkoxy groups, mixtures of hydroxyl andalkoxy groups, chlorine, etc. Examples of supports may include silica,silica-alumina, titanium oxide, zirconium oxide, or zeolitic materials,such as MCM-41, MCM-48, SBA-15, etc., and composite materials includinga binder and a zeolite.

Alternative supports may include carbon and/or carbonaceous materials.Carbon and carbonaceous supports may have surface functional groups,such as hydroxyl, carbonyl, or both, to immobilize organometalliccompounds on the surface, as discussed earlier. To prepare supportedmetal oxide, hydroxide, or oxyhydroxide catalysts, the surfacefunctional groups may not be necessary, although may be useful in someembodiments. Carbonaceous supports may be prepared by controlled thermaldehydration of carbohydrates, such as wood, coconut shells, starches,cellulose, a mixture of starch and cellulose, sugar, methyl cellulose,and the like, at elevated temperatures. Carbonaceous supports may beeither unsupported or supported. To prepare supported carbonaceousmaterial, carbohydrates may be deposited on a suitable porous supportfollowed by controlled thermal dehydration at an elevated temperature,such as a temperature in the range from about 250° C. to 1000° C., in aninert atmosphere or an atmosphere composed of an inert gas and a smallamount of oxygen, steam or a mixture thereof. Supports for carbonaceousmaterials may include any inorganic materials, such as alumina, titania,silica, zirconia, synthetic and natural clays, includingsilica-aluminas, and other supports as known in the art.

The supports, in some embodiments, may require removal of condensedwater in the pores prior to contacting organometallic compounds with thesupports to perform immobilization. Condensed water on a support isdefined here as water content that may be removed by drying the supportat a temperature in the range from about 50° C. to about 400° C. in drygas flow or under a vacuum, depending upon chemical composition of thesupport. Solid catalysts used herein may be prepared by immobilizing oneor two organometallic compounds or reacting one or multiple solublemetal compounds with surface functional groups having active catalystsites on a porous solid support. Immobilization may be performed, forexample, by using techniques such as grafting, tethering, adsorption,etc. For example, catalyst preparation techniques for organometalliccompounds such as titanium alkoxides on porous supports has beendisclosed in the '672 publication.

A second type of catalyst useful in embodiments disclosed hereinincludes a metal oxide, mixed metal oxides, or oxyhydroxides depositedon a porous support. Examples of this type of catalyst are alsodisclosed in the '672 publication.

Supports may be in the form of pellets, extrudates, spheres, granules,honey comb, and the like, in sizes ranging from about 1 mm to about 5 mmfor various fixed bed reactors. Alternatively one may choose to usewoven cloth or mesh made out of fiberglass or carbon fiber or both assupport along with structured packing materials, which are suitablyshaped and sized properly depending on type of reactors. Supports inpowder or microsphere forms may also be used for the preparation ofcatalysts to be used for slurry or stirred reactor.

Preparation of the second type of catalysts described above may notrequire a support having surface hydroxyl group. However, surfacefunctional group-containing supports, such as silica, carbonaceousmaterial, alumina, etc., may also be used to prepare metalhydroxide/oxide catalyst by grafting metal alkoxides, such as titaniumalkoxide, on a silica, followed by steaming or hydrolyzing and/or dryingat a temperature from about 90° C. to about 500° C.

Another method for preparing metal oxide or oxyhydroxide catalystsincludes depositing a salt of a desired element or a mixture of salts oftwo different elements on a support followed by calcining at atemperature from 300° C. to 1000° C. to decompose the salts to metaloxides.

Under process conditions, transesterification and disproportionation ina catalytic reaction zone may occur simultaneously as the concentrationof an alkyl aryl carbonate in the reaction medium increases. The twocauses for the catalyst deactivation discussed above, leaching andpolymer deposition, also occur simultaneously under reaction condition.While polymer deposition does not cause permanent damage to a catalyst,the active metal component leaching out of heterogeneous catalysts underreaction conditions does result in permanent damage to the catalyst. Atlow conversion levels for the transesterification or at lowconcentrations of alkyl aryl and diaryl carbonates, catalystdeactivation is mostly caused by dissolution of active metal catalystcomponents from a solid catalyst into the reaction medium. In otherwords, the cause of permanent catalyst deactivation under all reactionconditions is metal leaching.

As the transesterification conversion increases, polymer deposition onthe catalyst causes even faster catalyst deactivation. Polymerdeposition is primarily the result of undesired side reactions of alkylaryl and diaryl carbonates (and potentially trace amounts of polyhydroxyl aromatic compound impurities in a phenol feed and produced byundesired side-reactions in minute quantities). Therefore, tocontinuously produce diphenyl carbonate from phenol and a dialkylcarbonate, such as diethyl carbonate or dimethyl carbonate, in thepresence of heterogeneous catalysts, it may be required to addresscatalyst deactivations caused by both (1) polymer deposition and (2)dissolution/leaching of active metal catalyst components. Polymerdeposition may be addressed by controlling conversion, concentration ofaromatic carbonates, or both in the catalytic reaction zone, asmentioned above, and via catalyst reactivation, such as disclosed in the'672 publication. Leaching is addressed via adding a trace amount ofsoluble organometallic compound, as described above.

Immobilizing (e.g., grafting, tethering, adsorption, etc.)organometallic compounds or soluble metallic compounds on a support suchas silica or carbonaceous material for alcoholysis and/ortransesterification of a dialkyl carbonate with phenol may be carriedout in a single reaction zone step or multiple reaction zone steps.Examples of the organometallic compounds disclosed include metalalkoxides, alkoxy chlorides, oxyalkoxides, carboxylates, carbonates,etc., of Group II, III, IV, V and VI elements. Examples of active metalsinclude Mg, Ca, Zn, La, Ac, Ti, Zr, Hf, V, Nb, Ta, Cr, Mo, W, Sn, Pb,Sb, etc. In various embodiments, tin alkoxides, alkyl tin alkoxides,alkyl tin oxides, alkyl tin hydroxides, dialkyl tin dichloride, alkyltin trichloride and a mixtures of these species as well as metaloxyalkoxides [(RO)_(n)MO] and metal alkoxy hydroxide [(RO)_(n)M(OH)_(x)]or oligomers of these oxyalkoxides and alkoxy hydroxide are included,where M is a Group IV, V or VI element, n=2, 3 or 4, x=0, 1, 2 or 3, andn+x=4, 5 or 6. In selected embodiments, the organometallic compound maybe one or more of a titanium alkoxide or phenoxide, alkyl aryl titanate,or titanium salt of a carbonic acid monoester. It should be understoodthat metal alkoxides include monomers, various oligomers, or a mixtureof various monomer and oligomer species, depending on the carbon chainlength and structure of the alkyl group of an alkoxide or aryloxide[see, for example, Coordin. Chem. Rev., 2 (1967) 299-318; J. Chem. Soc.,3977 (1955)].

As described herein, alkoxides of a transition metal include all thespecies of monomer and various oligomers. For example, while titaniumethoxide [Ti(OEt)₄] exists mostly as trimer in boiling ethanol orbenzene, sterically hindered titanium alkoxides, such as titaniumisoproxide, are monomeric in boiling hydrocarbon solutions. For example,titanium isopropoxide is believed to exist mostly as monomer in aboiling toluene solution.

Porous supports used in various embodiments disclosed herein may havesurface hydroxyl groups, alkoxy groups, or both. To prepare the poroussupport, porous metal oxide supports, such as titanium oxide, zirconiumoxide, molybdenum oxide, and vanadium oxide, may be treated with astream containing one or more of an alcohol, an organic carbonate, suchas dimethyl carbonate, diethyl carbonate, etc., at a temperature in therange from about 130° C. to about 400° C. in some embodiments, from 150°C. to 350° C. in other embodiments, in a vapor phase, liquid phase, or avapor-liquid system. The stream may contain water from 0 wt % to about20 wt % in some embodiments; from 0 wt % to about 10 wt % in otherembodiments; and from about 0.005 wt % to about 5 wt % in yet otherembodiments. As water has little solubility in DMC and DEC, the streammay contain appropriate amounts of methanol and/or ethanol as solventfor water. Commercially available silica gel or silica, having surfacehydroxyl groups, may be used in some embodiments. Optionally one mayperform treatment of silica with liquid water, steam, or a mixturethereof at a temperature from about 80° C. to about 500° C. followed bydrying at a temperature from about 70° to about 800° C. in someembodiments, and from about 80° C. to about 500° C. in otherembodiments.

Silooxane and siloxane compounds of the transition metals may also beused to prepare solid catalysts immobilized on porous supports or metaloxide catalysts as described above. Examples of silooxane and siloxanecompounds are (RO)_(n−x)M[-O—Si(O—R)₃]_(x), M(O—SiR₃)_(n),(R₃SiO)_(n−2)MO, etc., where each R is independently an alkyl or arylgroup, n=3, 4, or 5, x=1 or 2, n+x=4, 5 or 6, and M is a transitionmetal of Group IV, V or VI as described above. Other silicon-metalcompounds are within the scope of embodiments disclosed herein, as longas the immobilization results in catalytic activity of the solidcatalysts. Silooxane and siloxane compounds of the transition metals mayalso be used as soluble organometallic compounds in the processarrangements disclosed in the '672 and the '214 publications, as well asreactive distillation column reactors. Various oligomeric and polymericheterosilooxane or heterosiloxanes of the transition metals may also beused, may be used to prepare immobilized solid catalysts, or may be usedas soluble organometallic compounds in various embodiments. As describedabove, disproportionation of EPC or MPC to DPC and DEC or DMC may beperformed in the absence of a solid catalyst in the second reactionzone, and useful active catalytic species include silooxane or siloxanecompounds of a transition metal such as Ti.

The metal oxides and alkoxides of siloxanes may include variousoligomers. Various oligomers may be found in the publications by Bradely[D. C. Bradley, Coordin. Chem. Rev., 2 (1967) p.p. 299-318); J. Chem.Soc., (1955) 3977]. One may choose to add a trace amount of one of thesecompounds into the feed to the reaction zone to obtain stable catalystactivity.

When carrying out disproportionation of alkyl aryl carbonates to producediaryl carbonate and dialkyl carbonate in the second reaction zone inthe presence of a homogeneous catalyst, the homogeneous catalyst may bea mixture of alkyl aryl titanates, titanium salts of carbonic acidmonoesters, and siloxane compounds of titanium discussed above. It isunderstood that the homogeneous catalysts may originate from the solidcatalysts and soluble catalysts used in the transesterification reactionzone.

As various organo-metallic compounds disclosed herein are sensitive tomoisture in the feed stream, it is important to control the watercontent in the feed stream to the reaction zone. In some embodiments,the moisture content of the feed stream is less than about 700 ppm; lessthan about 600 ppm in other embodiments.

The solid metal alkoxide catalyst immobilized on a support may beprepared by in situ techniques inside the reactor or may be preparedoutside reactor. For in situ preparation, a predetermined amount of asuitable support is placed in a reactor, followed by drying at a propertemperature to remove at least a portion of any condensed water. Thesupport is then contacted with a solution containing soluble metalalkoxide or mixed metal alkoxides of a transition metal or metals at atemperature in the range from about ambient to about 260° C. (500° F.)in some embodiments, and from about 37° C. to about 204° C. (about 100°F. to about 400° F.) in other embodiments. Contacting may be performedfor a period of time from about 5 minutes to about 24 hours in someembodiments, and from about 15 minutes to about 15 hours in otherembodiments, and may depend on the temperature and the concentration ofactive metal component in solution. After draining out excess metalalkoxide solution from the reactor, the catalyst in the reactor may bewashed with a solvent (usually the same solvent used to prepare themetal alkoxide solution) prior to use in disproportionation ortransesterification reactions. The solvent may be an alcohol, an ether,a hydrocarbon, a mixture of hydrocarbons and an alcohol, or a mixture ofa dialkyl carbonate and phenol or an alcohol, or mixtures of all ofthese.

Alternatively, metal oxide, mixed-metal oxide, or metal hydroxidecatalysts, where the metal is one or more from Groups II, III, IV, V andVI of the Periodic Table, may also be used according to embodimentsdisclosed herein. Some metal oxide catalysts are known in the art. Forexample, according to P. Iengo et al., Appl Catal. A: General 178 (1999)97-109, titanium oxide catalysts supported on silica may be prepared bygrafting titanium isoperoxide and then steaming/calcination, thesupported catalyst having a strongly modified original silica surfaceresulting in a catalyst different from those obtained by impregnationand co-precipitation.

To prepare supported metal or mixed metal hydroxide or oxyhydroxidecatalysts, one may hydrolyze the grafted metal alkoxide catalysts, asdescribed above, followed by drying at a temperature in the range fromabout 50° C. to about 110° C. In some embodiments, drying may not benecessary.

Preconditioning of unsupported metal oxide catalysts may be performedprior to performing reactions producing organic carbonates. Thepreconditioning is performed by contacting a porous metal oxidecatalysts, such as titanium oxide, zirconium oxide, molybdenum oxide, orvanadium oxide, with a stream containing an organic carbonate, such asdimethyl carbonate, diethyl carbonate, etc., at a temperature in therange from about 125° C. to about 450° C. in some embodiments, and fromabout 150° C. to about 350° C. in other embodiments, where the organiccarbonate may be in the vapor phase, liquid phase, or a mixed phase.Preconditioning may be performed for a period from about 2 minutes toabout 50 hours in some embodiments, and from about 4 minutes to about 24hours in other embodiments. The stream containing organic carbonate mayinclude water and an alcohol, where the water may be present fromgreater than 0 wt % to about 10 wt % in some embodiments, and from about0.005 wt % to about 4 wt % in other embodiments. Selectivity of thecatalyst may be improved by the preconditioning. After preconditioning,the metal oxide catalysts may be dried at a temperature from about 80°C. to about 300° C. in an inert gas flow for a period of time from about2 minutes to about 6 hours.

Two types of mixed metal oxide catalysts may be used for thetransesterification of a cyclic carbonate with an alcohol. The firsttype of mixed metal oxide catalysts may include one or more elementsfrom Groups III, IV, V, and VI of the Periodic Table supported on asupport. The second type of the mixed oxides includes a solid basecatalyst that contains one or two elements from Group II of the PeriodicTable and lanthanides or actinides on a support. Optionally, one may usea quaternary ammonium hydroxide grafted or tethered on a silica support.The oxide catalysts are usually supported on alumina or silica orprepared in the form of a mixed oxide or a solid solution. The elementsuseful for the second type of solid catalysts may include Mg, Ca, Zn,La, etc.

Active metal components of the second type catalysts may also leach outunder transesterification reaction conditions, resulting in catalystdeactivation. In fact, it has been found that silica supports may alsoleach out, only at a much slower rate than Group II active metalcomponents. Since alkali metal impurities on a silica support mayincrease the dissolution of silica into the reaction medium, minimalalkali metal impurites in silica support is highly desirable. By addingtrace amount of soluble organometallic compounds into a feed stream, thecycle length of a solid catalyst may be extended for fixed bed reactors.Examples of such soluble compounds include zinc 2-methoxyethoxide,calcium 2-methoxyethoxide, zinc 2-methoxypropoxide, zinc ethoxide, zincalkoxy alkyl carbonate, calcium 2-methoxyproxide, calcium ethoxide,calcium alkoxy alkyl carbonate, magnesium 2-methoxyethoxide, magnesium2-methoxyproxide, magnesium ethoxide, magnesium butoxide, magnesiumalkoxy alkyl carbonate, lanthanum alkoxide, lanthanum alkoxy alkylcarbonate, and Mg, Ca, and Zn propylene glycerides or glycolates, amongothers. A mixture of these may also be used.

Soluble compounds of Ca, Mg, Zn, and La may be obtained by reactingoxides or hydroxides of these metals with an alcohol, an organiccarbonate, or a mixture of an organic carbonate and an alcohol attemperature from about 105° C. (221° F.) to about 260° C. (500° F.) insome embodiments, from about 149° C. (300° F.) to about 227° C. (440°F.) in other embodiments, in liquid phase or mixed phase (liquid andvapor) system. Solutions prepared in this manner may be useful foradding trace amounts of these metals into the feed stream to a reactorso as to obtain long cycle time. Total amount of active metal or metalcomponents on a solid metal alkoxide, metal hydroxide, or metal oxidecatalyst may range from about 0.02 wt % to about 20 wt %, preferablyfrom about 0.05 wt % to about 12 wt %.

Improved Catalyst Cycle Length and Service Life

The solid catalysts disclosed herein may have long cycle lengths and maybe able to undergo catalyst regeneration and reactivation many times,resulting in long catalyst service times. The techniques of extendingcatalyst cycle length and catalyst reactivation disclosed herein makecatalysts typically interesting for laboratory purposes useful incommercial production of various organic carbonates. It is surmised thatwhether starting with supported metal oxide catalysts or metal alkoxidecatalysts immobilized on silica, at steady state, the active catalystsare organometallic compound species immobilized on silica. To illustratethe benefit of adding a trace amount of active metals in the feed,various experiments were performed, and will be described in more detailbelow. Briefly, in one experiment, a titanium oxide catalyst (6 wt % Ti)supported on a silica gel, was regenerated by depolymerization afterserving about 350 hours, and recovered less than half of its originalactivity of the transesterification of diethyl carbonate with phenol. Itwas found that more than half of the Ti had been leached out of thecatalyst into the reaction medium during the service time. In anotherexperiment, a titanium butoxide catalyst (4 wt % Ti) grafted on a silicagel lost over 90% Ti after serving only 171 hours for thedisproportionation of ethyl phenyl carbonate. After serving 173 hours,another titanium oxide catalyst (5.7 wt %) supported on silica gel, usedfor transesterification of propylene carbonate with ethanol to producediethyl carbonate and propylene glycol, lost 35% of the Ti on thecatalyst. From these findings, it is very clear that both supportedtitanium oxide catalyst and grafted titanium alkoxide catalyst areunsuitable for commercial reactors for the continuous production oforganic carbonates, such as dialkyl carbonate, alkyl phenyl carbonate,and diary carbonate, due to permanent catalyst deactivation in shortservice time. It appears that organic carbonates and/or reactionmixtures are reactive enough to cause slow formation of solubleorganometallic compounds in the reaction medium by reacting with solidcatalysts.

It has also been documented that DMC and DEC vapor streams may reactwith silica or titanium oxide to form tetra-alkyl orthosilica andtitanium tetra-alkoxide at a temperature higher than about 350° C. Thereactions of DMC and DEC with silica become easier in the presence of acatalytic amount of alkali metal on silica. Therefore, it was necessaryto discover a technique for simple catalyst reactivation and a method tomaintain constant surface concentrations of active metal components onthe catalyst for a sufficiently long cycle time to be acceptable forcommercial reactors.

Catalyst regeneration via depolymerization and metal redepositionaddresses problems associated with polymer deposition. However,regeneration via depolymerization fails to address the problemsassociated with continuous leaching of active metals from heterogeneouscatalysts under reaction conditions. The continuous loss of active metalfrom heterogeneous catalysts must be addressed to obtain long catalystcycle length suitable for commercial scale reactors. It was discoveredthat the effect of metal leaching from heterogeneous catalysts could beneutralized by adding trace amounts of active metal compounds into thefeed stream(s) to the first reactor in the case of a series of multiplereactor system. By adding a trace amount of soluble active metalcompound, metal leaching and redeposition are balanced or nearlybalanced, effectively maintaining a constant number of active sites onthe solid catalysts, resulting in a steady catalyst activity for a longcatalyst cycle time. It is understood that the soluble metal componentleaching out of the solid catalyst in the first reactor is a mixture ofvarious metal compound species. The metal compound species in themixture are not necessarily identical to the metal compound speciescoming into the first reactor. The metal leach and redeposition in thesecond reactor is also balanced in similar fashion. For parallelmultiple reactor systems, the addition of a trace amount of active metalcompound to all feed streams to the first primary reaction zone may berequired. Therefore, catalyst reactivation (in situdepolymerization/surface conditioning and metal redeposition) and addinga trace amount of active metal compound followed by metal redepositionmay address both metal leaching and polymer deposition. Catalystreactivation may be performed in two steps: (1)depolymerization/conditioning of the catalyst surface, and (2)redeposition of active metal components. The catalyst surfaceconditioning is necessary to immobilize titanium alkoxide on silicasupport surface. A fresh immobilized catalyst or a reactivated catalystloses catalytic activity continuously, resulting in unacceptably shortcycle time that is unsuitable for large commercial reactors. The loss ofan original catalyst activity to about half activity fortransesterification of a dialkyl carbonate with phenol takes about80-150 hours-on-stream time, which is clearly inadequate for continuousoperation of commercial reactor. By adding trace amounts of solubleactive metal compound and performing the two-step reactivation, it isnow possible to extend catalyst cycle length and carry out catalystreactivation multiple times for the continuous production of variousorganic carbonates.

Depolymerization of a deactivated catalyst may be performed bycontacting the catalyst with a stream containing a hydroxy compound or amixture of hydroxy compounds in situ at a temperature from 102° C. (215°F.) to 316° C. (600° F.) in some embodiments, from 104° C. (220° F.) to232° C. (450° F.) in other embodiments, for a period of time from about10 minutes to about 50 hours in some embodiments, and from 30 minutes to15 hours in other embodiments. The depolymerization may be carried outin vapor phase, in liquid phase, a mixed phase, or in liquid phasefollowed by in vapor phase, or in reverse order. Depolymerizationproducts may include phenol, alcohol, carbon dioxide, multihydroxybenzene, dialkyl carbonate, alkyl phenyl carbonate, and heaviercompounds.

Examples of hydroxy compounds to be used for depolymerization on acatalyst are alcohols (preferably methanol or ethanol), water, or amixture thereof. If dimethyl carbonate is used as one of the feedstocksto produce methyl phenyl carbonate and diphenyl carbonate, methanol or amixture of water and methanol may be used for the depolymerization. Ifdiethyl carbonate is used as one of the feedstocks, ethanol or a mixtureof water and ethanol may be used for the depolymerization. One may alsouse a mixture of methanol and ethanol. When water is used in thedepolymerization, the water content in the mixture may be in the rangefrom greater than zero wt % to less than 100 wt % in some embodiments;from 10 ppm by weight to 15 wt % in other embodiments; and from 15 ppmby weight to 5 wt % in yet other embodiments. An azeotropic mixture ofwater (4 wt %) and ethanol is very effective for the depolymerizationwhere diethyl carbonate is used as one of the feed stocks. In someembodiments, a mixture of water and alcohol may be preferred over eitherwater or alcohol alone, allowing for the conditioning of the catalystsurface for the redeposition of active metal component. Additionally, amixture of water and an alcohol may be more effective for thedepolymerization and surface conditioning than alcohol or water alone.

One may also use a solvent in the depolymerization. Useful solvents mayinclude benzene, toluene, xylenes, pentane, hexane, octane, decane,tetrahydrofuran, ether, etc. or any mixtures of such solvents. Theconcentration of solvent in a depolymerization mixture may range from 0wt % to about 90 wt %.

Depolymerized catalyst may be dried to remove excess water on thecatalyst and to control the population of surface hydroxyl groups priorto redeposition of active metal component on the catalyst. The in situdrying may be carried out at a temperature from about 49° C. (120° F.)to about 427° C. (800° F.) in some embodiments, from 65° C. (150° F.) to316° C. (600° F.) in other embodiments, in an inert gas flow, for aperiod of from about 15 minutes to 40 hours under ambient pressure or asubatmospheric pressure, prior to performing redeposition of activecatalyst component. Improper catalyst surface preconditioning may resultin only partial recovery of catalyst activity. The depolymerizationtechnique disclosed herein may be used in any process for the productionof aromatic carbonates or any reaction where organic carbonates areinvolved as reactants, products, or both.

The depolymerization technique disclosed may also be useful forreactions producing organic carbonates in the presence of a homogeneouscatalyst. For the regeneration of a homogeneous catalyst system, analcohol solution must be fairly dry so that water content may not exceedabout 0.01 wt %. Therefore, the catalyst regeneration techniquedisclosed herein may be useful for any process for producing organiccarbonates.

The effluent stream from a reactor during depolymerization may contain atrace amount of active metal component, depending on how thedepolymerization is performed. This stream may also contain phenol, DEC,small amounts of phenetole, EPC and heavier compounds as majordepolymerization products. If desired, one may attempt to recover theuseful components such as phenol, ethanol, alkyl phenyl carbonate andDEC from this stream.

The redeposition of active metal component on a depolymerized andsurface conditioned support may be carried out in a similar manner toimmobilizing a metal alkoxide on a support as described above. Theimmobilizing of a metal alkoxide on a support may be carried out in asingle step or in multiple steps. A reactor containing reactivatedcatalyst is then ready to go back into service.

Addition of trace amounts of soluble active metal components ofcatalysts into the feed stream, as described above, may result in astable catalyst performance for prolonged cycle time. As an example,transesterification of DEC with phenol was performed in an up-flow,once-through, fixed bed reactor with about 45 to about 60 ppm by weightTi added to the feed stream. There has been little sign of catalystdeactivation during more than 14 months of continuous on-stream time.

The addition of a trace amount of active metal compounds as disclosedherein may be useful for continuous commercial production of variousorganic carbonates or carbamates. The reactions producing organiccarbonates may be performed in a single reactor, a series of multiplereactors, or a multiple parallel reactor system, as a specific reactionsystem dictates. For example, the reactions may be carried out in asingle catalytic distillation column reactor or a series of multiplecatalytic distillation column reactors, where a solid catalyst or twodifferent solid catalysts are placed. Optionally a series of multipleslurry reactors may also be used to produce an organic carbonate. Theaddition of a trace amount of soluble active metal component to a feedstream may be to only the first reactor in a series of multiplereactors. The desirable amount of trace amount of active metal componentin a feed stream depends on the specific active metal element forspecific feed components. For the transesterification of a dialkylcarbonate with phenol, it may range from about 15 ppm to about 400 ppmby weight in some embodiments; from about 20 ppm to about 300 ppm inother embodiments; and from about 25 ppm to about 200 ppm in yet otherembodiments, depending on the metal. For a feed stream composed ofdiethyl carbonate and phenol, for example, a desirable amount of Ti maybe from about 20 ppm to about 150 ppm in some embodiments; and from 30ppm to 100 ppm by weight in yet other embodiments. The amount of theactive metal component in the feed stream is approximately one or twoorder of magnitude lower than the concentration of homogeneous catalystin the reaction medium of prior arts.

The Ti concentration in a reactor effluent stream is usually in a rangefrom about 20 ppm to about 100 ppm, depending on the amount of activemetal concentration in the feed stream to a reactor. At this level, itis generally not economically favorable to recover Ti from the reactoreffluent streams for recycle, although one may choose to do so. Theactive metal component in the reactor effluent stream may be recoveredfrom heavy bottom streams of the crude DPC recovery column as solidmaterial and converted to soluble organometallic compound to be reusedby reacting with organic carbonate or a mixture of organic carbonate andalcohol at an elevated temperature. The recovered organometalliccompound may be a metal alkoxide, a metal alkoxy alkyl carbonate (metalsalt of a carbonic acid monoester) or a mixture of these.

To recover the soluble active metal component in the bottom stream ofthe DPC recovery column as solid material, the heavy waste bottom streamfrom DPC recovery column may be treated with hot water or a mixture ofsteam and water to precipitate the metal component as a solid. In caseof a solid titanium containing catalysts, the solid titanium precipitatein the aqueous phase is separated from liquid by using conventionalmethods, such as filtration or centrifuge. The separated solid isconverted to soluble material by treating with a liquid streamcontaining a dialkyl carbonate or a mixture of dialkyl carbonate and analcohol at a temperature from 121 to 343° C. (250 to 650° F.) underpressure for a period from 10 minutes to 80 hours in some embodiments,and from 20 minutes to 45 hrs in other embodiments. The pressure issufficiently high so that dialkyl carbonate or a mixture of alcohol anddialkyl carbonate should, at least partially, exist as liquid in areaction vessel, such as an autoclave, tubular reactors, distillationcolumn reactor system, etc. Optionally, the liquid stream may contain aninert solvent such as benzene, toluene, hexane, heptane, ether, etc.Examples of the liquid stream are mixtures of ethanol and DEC ormethanol and DMC. The content of dialkyl carbonate in a mixture of analcohol and dialkyl carbonate may range from 0.1 wt % to less than 100wt %.

The reactions producing organic carbonates or carbamates may beperformed in a single reactor or in a series of multiple reactors invarious arrangements of the reactors with suitably arranged distillationcolumns for the cost effective separation of reaction products and forthe recycle of unreacted reactants. Alternatively the reactions may beperformed in a single or a multiple parallel reactors. Various otherarrangements of reactors and distillation column may be devised by thoseskilled in the art.

A reaction may be performed in a single catalytic distillation column,in a series of multiple catalytic distillation columns, in a series ofmultiple fixed tubular or tank reactors or any combination of differenttypes of reactors. When three catalytic distillation columns are used toproduce DPC, a solid catalyst is placed in the first two reactors inseries for the transesterification. The third distillation columnreactor may contain a solid catalyst or alternatively may not contain asolid catalyst. The disproportionation in the third reactor may beperformed by utilizing only soluble homogeneous catalyst present in thereaction medium.

Reactions performed in traditional tubular fixed bed reactors may beperformed in up-flow or down-flow mode. The reactions producing alkylaryl carbonate, such as EPC, and diaryl carbonate, such as DPC, forexample, may be performed in a liquid phase, but may also be performedin a mixed phase system in the presence of one or more solid catalysts.The two reactors in series for the transesterification may alternateperiodically between being first and second reactor to prolong cycletime. The third reactor for disproportionation of EPC to produce DPC andDEC may be performed in the lower half of the phenol recovery column,which may be operated at sub-atmospheric pressures. In selectedembodiments, processes disclosed herein may be useful for producingdiphenyl carbonate by performing transesterification of diethylcarbonate with phenol followed by disproportionation of ethyl phenylcarbonate.

A trace amount of a soluble active metal compound, such as ethyl phenyltitanate or ethoxy titanium ethyl carbonate, or a mixture of titaniumalkoxide and alkoxy titanium alkyl carbonate, for example, may be addedinto the liquid reaction medium fed into the first reaction zone.Alternatively, disproportionation may be performed in a catalyticdistillation column to which the bottoms stream from a phenol recoverycolumn is introduced at a proper point in the upper mid-section of thecolumn. Disproportionation may also be performed in a catalyticdistillation column to which the bottoms stream from atransesterification reactor is directly introduced without phenolremoval (phenol may be recovered from the bottoms stream from the EPCdisproportionation column). The first reaction zone may include twocatalytic distillation columns in series or two parallel catalyticdistillation columns for transesterification of DEC with phenol. Thesecond reaction zone may include a catalytic distillation column reactorfor disproportionation of EPC to DPC and DEC. One may choose DMC inplace of DEC and MPC in place of EPC. Catalytic distillation columnreactors for the first reaction zone may be loaded with one or moresolid catalysts, such as titanium alkoxide immobilized on a silicasupport or titanium oxide supported on a silica support. In general,there are two alternative processes for the continuous production ofdiphenyl carbonate that may be used, in the case where more than tworeactors are used.

In a first process for the continuous production of diphenyl carbonate,there may be from three to seven catalytic distillation column reactorsin various embodiments, from three to four catalytic distillationcolumns in selected embodiments. Out of these catalytic distillationcolumn reactors, one or more may serve as a spare reactors to replacethe least active reactor out of the multiple reactors in service. Of themultiple distillation column reactors, two to six reactors may be usedto primarily produce EPC. The remaining catalytic distillation columnreactors may serve as a second reaction zone, wherein primarily the EPCdisproportionation to DPC and DEC occurs. DEC and at least a portion ofthe phenol in the stream from the first reaction zone are removed fromthe heavy effluent stream of the first reaction zone prior to enteringthe second reaction zone. Alternatively, removal of phenol in the heavyeffluent stream of the first reaction zone may be delayed until afterdisproportionation, depending on the concentration of phenol in thatstream. As the catalyst in service ages, the catalyst activity slowlydeactivates. There are three different options for rotating the multiplereactors in series between service for production of aromatic carbonatesand catalyst reactivation:

-   -   (1) cyclic rotation of all the reactors in sequential order        after a given service time with the oldest reactor coming out of        service for catalyst reactivation while bringing a reactor        having fresh or reactivated catalyst into service as the first        reactor in the series of multiple reactors (i.e., new→first        reactor, first reactor→second; second→third; and        third→reactivation or catalyst replacement), or optionally bring        a new reactor into service as the last reactor in the series, as        the second reactor moves up as the first reactor (reverse of        forward sequence presented);    -   (2) the reactors are divided into two groups of first reaction        zone and second reaction zone reactors with each group having a        spare reactor for rotation between service and catalyst        replacement/reactivation;    -   (3) bring the least active reactor in a series of the multiple        reactors out of service for the catalyst reactivation, as        necessary, and bring a spare reactor (wherein the catalyst        already had been reactivated) into service to replace the        reactor taken out of service.

In an alternative process, two reactors in series are used as the firstreaction zone. The sequence of the two reactors are periodicallyalternated between being the first reactor and the second reactor afterservicing for a given period of time, say every 6000 hours; thisrotation repeats as many times as necessary. There is no spare reactorfor the second reaction zone. This type of operation is possible due tothe addition of trace amounts of active metal compounds to the firstreactor in the series. DEC and phenol in the stream from the firstreaction zone are removed by distillation, and then the remaining streamis subjected to disproportionation of EPC to produce DPC in the secondreaction zone. There are two ways carrying out the disproportionation.

-   -   (1) In the first method, disproptionation is performed in the        presence of solid catalyst in a fixed bed reactor such as        catalytic distillation reactor. There is a spare reactor for the        replacement of the reactor in service. The deactivated catalyst        is subjected to catalyst reactivation described earlier.    -   (2) In the second method, the disproportionation is performed in        a catalytic distillation reactor in the absence of a solid        catalyst and there is no spare reactor. The active soluble metal        species in the stream coming from the first reaction zone serve        as a homogeneous catalyst for the disproportionation reaction.

It is understood that the catalytic distillation column, wherein a solidcatalyst is either present or absent, for the second reaction zone isdesigned such that the top half section of the column (phenol recoverysection) serves primarily to distill off phenol in the incoming reactionmixture from the first reaction zone and the bottom half section servesprimarily to perform disproportionation of EPC or MPC. In an alternativeprocess design, a phenol recovery column and a catalytic distillationcolumn are separated into two columns, although some disproportion mayoccur in the bottom section of phenol recovery column. As stated above,depending on the concentration of phenol in the incoming feed stream, aphenol recovery may be delayed until after disproportionation, althoughsome of phenol may be stripped off in the catalytic distillation columnas overhead vapor stream along with DEC. The catalytic distillationcolumn for disproportionation may be operated under a subatmosphericpressure.

FIG. 1 is a simplified flow diagram illustrating a process for thecontinuous production of DPC with three catalytic distillation columnsaccording to embodiments disclosed herein. Two catalytic distillationcolumns in series serve as the first reaction zone fortransesterification of DEC with phenol to produce EPC and ethanol in thepresence of a solid catalyst, and a catalytic distillation column servesas the second reaction zone for disproportionation of EPC to produce DPCand DEC.

Referring now to FIG. 1, a process for the production of DPC from DECand phenol according to embodiments disclosed herein is illustrated. C1and C2 are catalytic distillation columns for performingtransesterification; C3 is an ethanol recovery column; C4 is a DECrecovery column (phenetole purge column); C5 is a catalytic distillationcolumn for disproportionation and phenol recovery; C6 is an EPC recoverycolumn; and C7 is a DPC recovery column

The columns C1 and C2 are a series of catalytic distillation columns,wherein structured packing devices are placed in reaction zone R1 andR2, respectively. The specially structured packing devices contain solidcatalyst. Phenol and DEC containing feed streams 1 and 4, respectively,are introduced to a tray in the upper section of catalytic distillationcolumns C1 and C2, above the catalytic reaction zones R1 and R2. Themole ratio of DEC to phenol in the fresh DEC and fresh phenol feedstreams may be approximately 1:2. The mole ratios of DEC to phenol inthe catalytic reaction zones R1 and R2, however, are controlled to befrom about 12:1 to about 1:2.5 in some embodiments; from about 10:1 toabout 1:2 in other embodiments; and from about 7:1 to about 1:1 in yetother embodiments.

A soluble organometallic compound is also introduced to a tray in thetop section of C1 via flow line 3. For example, for atitanium-containing solid catalyst in reaction zones R1 and R2, asolution containing a soluble titanium compound such asTi(Oet)_(4−x)(OPh)_(x) (where x is 0, 1, 2, 3 or 4), or titanium saltsof carbonic acid monoesters, such as ethoxy titanium ethyl carbonates,or mixtures of these, may be introduced to the top of the firstcatalytic distillation column reactor C1. The solvent for the catalystsolution can be DEC, a mixed solution of DEC and phenol, a mixedsolution of DEC and ethanol, or a mixed solution of DEC, ethanol, andphenol, for example.

The flow rate of the catalyst solution may be controlled such that theconcentration of titanium in the liquid stream above the catalyst in thefirst column reactor is from about 20 ppm to about 100 ppm active metal(titanium for the example catalyst solutions listed in the previousparagraph) by weight in some embodiments; from about 25 ppm to about 80ppm by weight in other embodiments; and from about 30 ppm to about 70ppm by weight in yet other embodiments.

Overhead vapor streams 6 and 14 from the catalytic distillation columnsC1 and C2 are sent to an ethanol recovery column C3 via flow line 8.This overhead stream may also contain minor amounts of by-products suchas diethyl ether and carbon dioxide and a trace amount of phenol.Diethyl ether and carbon dioxide may be removed as overhead vapor stream9. Ethanol may be recovered via a side drawn from column C3 via flowline 10. The bottom stream 11 may recycle DEC from column C3 tocatalytic distillation column reactors C1 and C2 via flow lines 12 and13, respectively.

Column C1 may be operated such that the temperature of the catalyticreaction zone R1 is in the range from about 160° C. to about 210° C.(about 320° F. to about 410° F.). The overhead pressure in column C1 maybe within the range from about 2 bar absolute to about 4.8 bar absolute(about 14.7 psig to about 55 psig). The bottom stream 7 from the firstcatalytic distillation column C1 may be introduced to the top ofcatalytic distillation column C2, which may be operated such that atemperature in the catalytic reaction zone may be in the range fromabout 162° C. to about 216° C. (about 325° F. to about 420° F.) and thecolumn may be operated at an overhead pressure in the range fromsub-atmospheric, about 1 bar (0 psig), to about 4.5 bar (51 psig).Optionally small fractions of recycle or fresh DEC stream can beintroduced to the column C1 and C2 via flow lines 4 a and 4 b,respectively.

The concentration of EPC increases moving down stages in the catalyticdistillation columns C1 and C2. As some disproportionation of EPC to DPCand DEC occurs in column reactors C1 and C2, the concentration of DPCincreases as well. The bottom stream 15 from distillation column reactorC2 is sent to DEC recovery column C4, where DEC is recovered in overheadvapor stream 16. Column C4 may be operated from a temperature of fromabout 127° C. to about 204° C. (about 260° F. to about 400° F.) at anoverhead pressure from about 0.3 bar (about 4 psia) to about 1.5 bar(about 22 psia). Stream 16 may be introduced to ethanol recovery columnC3 to separate DEC and phenol that may be in stream 16, where the DECand phenol may be recycled via lines 11, 12, 13 to C1 and C2.

The overhead stream 16 from column C4 may also contain DEC and smallamounts of phenetole, phenol, and ethanol. A side draw stream 18 fromcolumn C4 may be used as a phenetole purge stream, minimizing buildup ofphenetole in the system.

Bottoms stream 17 from column C4 contains homogeneous catalyst speciescoming from columns C1 and C2. Bottoms stream 17 may be introduced at asuitable position in the top section of distillation column C5. ColumnC5 may be used to perform disproportionation of EPC, and may optionallycontain a heterogeneous catalyst in reaction zone R3.

The column C5 may be designed and operated to serve two purposes: removephenol in stream 17 and the co-product DEC from the EPCdisproportionation as overhead stream 19; and, disproportionation of EPCto form DPC. Column C5 is operated so that the homogeneous catalyticreaction zone R3 temperature is in the range from about 165° C. to about210° C. (about 330° F. to about 410° F.), and the column overheadpressure is from about 0.07 bar (about 1 psia) to about 0.6 bar (9psia).

The overhead vapor stream 19 from C5, containing DEC and phenol, may berecycled to columns C1 and C2 via streams 20 and 21, respectively. C5bottoms stream 22 (which contains DPC, unconverted EPC, phenol,phenetole, heavies and soluble Ti catalyst) from C5 is introduced to EPCrecovery column C6, which may be operated at a temperature from about168° C. to about 213° C. (about 335° F. to about 415° F.) and undersub-atmospheric pressures in the range from about 0.03 bar (about 0.4psia) to about 0.55 bar (8 psia).

EPC Column C6 bottoms stream 25 is introduced to the DPC recovery columnC7 to recover DPC as side-draw stream 27. The DPC recovery column C7 isoperated under high vacuum (e.g., <0.03 bar (<0.4 psia). Overhead stream26 may be combined with EPC column C6 overhead stream 23 and recycled tocolumn C5 via line 24.

DPC recovery column C7 bottoms stream 28, containing heavies and solublecatalyst may be recovered or disposed. If desired, where titaniumcatalysts are used and fed to the reactor, for example, one may recovertitanium as soluble Ti catalyst (Ti(OEt)₄ or a mixture of Ti(OEt)₄ andethoxy titanium ethyl carbonates) for recycle, as discussed earlier. Asone disposal method, stream 28 can be sent to a titanium refinery torecover Ti. One can recover DPC from stream 22 in alternate recovery andpurification trains, which are within the knowledge of those of ordinaryskill in the art.

Alternatively, as mentioned above, the disproportionation of EPC to DPCand DEC may be performed in the presence of a solid catalyst in thecatalytic distillation column C5. Solid catalysts in C5 may, however,deactivate faster than solid catalysts in the second catalyticdistillation column C2, as illustrated in FIG. 1.

As discussed above, where C5 includes a solid catalyst, various optionsmay be used to cycle the distillation column reactors so as to maintainsufficient catalytic activity for the process. Sufficient valves andpiping, not illustrated, may be provided to allow for the cycling of thereactors, and is within the skill of those in the art.

FIG. 2 is an alternative process flow diagram, where like numeralsrepresent like parts. There are a similar number of catalyticdistillation columns to perform transesterification anddisproportionation, and columns for the material separations as inFIG. 1. However, a fraction of overhead stream 19 from catalyticdistillation column C5 may be recycled back to DEC recovery column C4via line 30. Recycle via line 30 may thus allow an alternative methodfor purging phenetole.

FIG. 3 illustrates another alternative process flow scheme according toembodiments disclosed herein, similar to FIGS. 1 and 2, where likenumerals represent like parts. The first catalytic distillation columnC1 is operated in more or less similar fashion to the previous cases(FIGS. 1 and 2). However, the operation of the second catalyticdistillation column C2 and the DEC recovery column C4 are operated indifferent fashions from previous cases. Column C2 may be operated at ahigher temperature and a lower pressure than previous cases. The recycleDEC stream 13 is introduced into the bottom of the column C2. Optionallya part of the fresh DEC stream 4 may also be introduced to the bottom ofcolumn C2. Due to operation at higher temperature and lower pressure, C2bottoms stream 15 contains less DEC. C2 overhead stream 14 containsphenetole, among other components. Stream 14 may be introduced to columnC4, where C4 bottoms stream 17 is the phenetole purge stream.

Extremely pure DPC may be prepared from crude DPC products producedaccording to embodiments disclosed herein. The high-purity DPC may beproduced by fractional crystallization using hydrocarbon-ether mixtures,such as hexane-diethyl ether mixtures. In some embodiments, the onlydetectable impurity, other than phenol, in the purified DPC product isxanthone in an amount up to about 0.5 ppm by weight. The phenol in thepurified DPC may be a trace amount, such as from about 5 to about 17 ppmby weight. Trace analyses of impurities in DPC produced according toembodiments disclosed herein indicate that the purity of the DPCobtained is much higher than for DPC that may be obtained from commonvendors for laboratory chemicals.

EXAMPLES

With the exception of Experiment 1, all transesterification reactions ofDEC with phenol were performed in up-flow boiling point reactor.Therefore, vapor and liquid phases coexists in the catalytic reactionzone. The dimension of the fixed bed reactor was 1.3 cm (½ inch)diameter by 6.5 cm (25 inches) long. The reactor had separatelycontrolled top and bottom heating zones. The fixed bed reactor wasmounted vertically. The volume of solid catalysts was 25 ml.

Comparative Experiment 1

Experiments for transesterification of DEC with phenol were performed inthe presence of homogeneous titanium alkoxide catalyst by using astirred 50 ml autoclave reactor. The autoclave was filled with about 35ml of a DEC/phenol mixture, as given in Table 1. The autoclave wasimmersed in an oil bath to control reaction temperature. Afterperforming the reaction, the autoclave was removed from oil bath andquenched with cold water. No diphenyl ether was observed in any reactionmixture. The results of the reactions are listed in Table 1.

TABLE 1 Catalyst DEC/ Selectivity Amount PhOH Phenol of EPC Temperature,Pressure, Duration Catalyst (wt. Mole Conversion and DPC ° C. (° F.) bar(psig) (h) Type ppm) Ratio (mole %) (mole %) 174 (345) 2.7 (25) 2Titanium 4767 2.32 4.54 62.97 Ethoxide 174 (345) 2.7 (25) 3 Titanium4767 2.32 15.45 15.24 Ethoxide 174 (345) 2.7 (25) 4 Titanium 4767 2.3214.93 20.45 Ethoxide 174 (345) 2.5 (22) 2 Titanium 42 2.35 0.13 91.2n-butoxide 174 (345) 2.7 (25) 4 Titanium 42 2.35 3.56 95.01 n-butoxide

When the Ti concentrations in the feed solutions were 4767 ppm byweight, the maximum phenol conversion after about 3 hours reaction timewas about 15% with very poor selectivity (<20.5 mole %) of EPC and DPC.When the reaction time was 2 hours, the conversion of phenol was lessthan 5%; the selectivity was better, but still poor (63 mole %). Whenthe catalyst concentration was reduced to 42 ppm Ti by weight, theselectivity was much improved, but conversion was poor.

Comparative Experiment 2

The objective of this experiment was to obtain experimental data ofhomogeneous catalyst as a reference to compare with the results ofexamples according to embodiments disclosed herein. There was no solidcatalyst in the reactor. The 25 ml space for a solid catalyst in thereactor was empty. A reaction mixture of 73.3 wt % DEC and 26.7 wt %Phenol (2.19 mole ratio of DEC/PhOH) having various amounts ofhomogeneous catalyst of Ti(Oet)_(4−x)(OPh)_(x) (x=˜2) were passedthrough the reactor up-flow under various reaction conditions from 0 to768 hours on stream and then from 1266 to 1362 hours on stream. Theconcentrations of Ti in the feed mixture ranged from 59 ppm to 709 ppmTi by weight, as shown FIG. 4. The flow rate was 0.5 ml/min for most ofrun time. The history of feed flow rate is listed in Table 2.

TABLE 2 Transesterification Time on Stream (h) 0-333 333-354 354-426426-1362 Feed flow rate (ml/min) 0.5 0.6 0.4 0.5

Ethanol in the composite transesterification products was distilled offand then DEC was added to adjust the DEC/PhOH mole ratio to 2.19 byadding DEC to prepare second transesterification feeds. The secondtransesterification feeds had about 3.4 wt % EPC, about 250 ppmphenetole and about 300 ppm DPC by weight, on average. The history ofthe homogenous catalyst concentrations in the second transesterificationfeeds is listed in Table 3.

TABLE 3 Time on Stream (h) 768-834 834-930 930-1026 1026-1098 1098-11941194-1266 Ti ppm by weight 188 113 174 197 300 220

Using these feed mixtures, the second transesterification was performedfrom 768 to 1266 hours on stream time. The phenol conversion in FIG. 4for this time period is the overall conversion through both 1st and 2ndtransesterification.

The range of the reaction temperatures was from 174° C. to 210° C. (345°F. to 410° F.) as indicated in FIG. 4. The range of the reactorpressures was from about 2.9 to about 5.5 bar (about 27 psig to about 65psig). All reactions were carried out under boiling condition.Therefore, the reaction was performed as a mixed phase system (vapor andliquid). The temperatures in FIG. 4 are reactor bottom temperatures.

When the catalyst concentration in the reaction mixture is higher than118 ppm Ti, there were adverse effects of the catalyst on the conversionof phenol. The cause of this adverse effect is not completelyunderstood, but may be the effect of two ethoxy groups on Ti catalyst.Also, when the Ti concentration in a feed was higher than about 300 ppm,there was line plugging problems in the reactor effluent line due toprecipitation of Ti catalyst. Therefore, a pair of in-line filter wasinstalled to address the line plugging problems. When the concentrationof Ti was 59 ppm, the temperature effect on the phenol conversion ismoderate, indicating low activation energy of transesterification of DECwith phenol. The highest phenol conversion for the firsttransesterification was about 11.3 mole % with 337 ppm Ti at 204° C.(400° F.) and 4.5 bar (50 psig). The highest phenol conversion for thesecond transesterification was about 14.5 mole % at 193° C. (380° F.)and 2.9 bar (27 psig) with 188 ppm Ti concentration. The experiment alsosuggests lower conversion for liquid phase reaction (410° C. and 7.9 bar(100 psig)), as expected.

Experiment 3

The objectives of this experiment were to demonstrate (1) the in situpreparation technique of titanium n-butoxide immobilized on a silica gelsupport, (2) technique for catalyst reactivation, and (3) theperformance of a dual phase fixed bed reactor for thetransesterification. The dual phase of vapor and liquid in the catalyticreaction zone was created by boiling the reaction mixture.

45.74 g of granular silica gel (+8 mesh) was treated with a sodiumhydroxide solution (7.5 g NaOH in 550 ml water) at about 42° C.temperature for 7 minutes with stirring at ambient temperature. Thesilica gel was washed, first with cold water and then with hot water(about 80° C.), to remove trace amounts of sodium on the silica. Theresulting treated silica gel was dried at 125° C. for 2 hours and thenat 300° C. for 2 hours under nitrogen purge. The dried silica gelsupport had 23 ppm Na by weight. The treated silica gel support had thefollowing properties: 291 m²/g BET, 1.052 cm³/g pore volume and 16.3 nmaverage pore diameter.

25 ml of the dried granular silica gel support (about 9.3 g) was loadedin the reactor. A titanium n-butoxide solution was prepared bydissolving 27 g of titanium n-butoxide in 500 ml dried toluene. Thetitanium n-butoxide solution was placed in a reservoir. Aftercirculating the titanium n-butoxide solution through the reactor up-flowat 15 ml/min and ambient temperature for 15 minutes, the reactor washeated to 168° C. (335° F.) at a pressure of about 5.5 bar (65 psig).The circulation was continued at 168° C. (335° F.) for 4.5 hours andthen the reactor was cooled. After draining excess solution from thereactor, the supported catalyst was washed with dry toluene up-flow at 4ml/min for 1.5 hours. The washed catalyst was dried at 168° C. (335° F.)in 350 cc/min nitrogen gas (up-flow) for 2 hours. The resulting insitu-prepared titanium n-butoxide catalyst grafted on silica gelgranular support was tested for transesterification of DEC with phenol.

1st Cycle Transesterification: Transesterification of DEC with phenolwas performed in the presence of the in situ-prepared solid catalyst ina fixed bed boiling point reactor. A mixture of DEC and phenol (25.57 wt% phenol and 74.43 wt % DEC; 2.32 mole ratio of DEC/PhOH) was passedthrough the solid catalyst bed up-flow at 168° C. (335° F.), 2.4 bar (20psig) and at a 0.2 ml/min feed flow rate. This test constituted the 1stcycle of transesterification and the results are illustrated in FIG. 5.The catalyst reached its maximum activity (12 mole % conversion ofphenol) at about 40 hours on stream. After about 80 hours on streamtime, the catalyst lost most of its activity. This deactivated catalystwas subjected to a 1st reactivation as follows.

1st Catalyst Reactivation: The catalyst reactivation comprises twosteps; catalyst depolymerization/surface conditioning and redepositionof active titanium metal on the catalyst. After draining the reactor,the catalyst was washed with dried toluene (300 ml) up-flow at ambienttemperature and then the toluene was drained from the reactor. 0.19 gtitanium n-butoxide was dissolved in 2 liters of an ethanol solutionprepared by mixing 400 ml ethanol and 1700 ml toluene. The titaniumsolution was passed through the reactor at 2.2 ml/min up-flow at 168° C.(335° F.) and 12 bar (160 psig) for 13.5 hours. After draining excesstitanium solution from the reactor, the catalyst was dried at 168° C.(335° F.) under ambient pressure in 200 cc/min nitrogen up-flow for 45minutes. A titanium n-butoxide solution (135 g titanium n-butoxide in 2liters toluene) was circulated through the reactor at 15 ml/min up-flowat room temperature for 20 minutes and then at 168° C. (335° F.) and10.7 bar (140 psig) for 4 hours. After cooling, excess solution wasdrained from the reactor. The catalyst was washed with 4 ml/min tolueneup-flow for 1.5 hours. The washed catalyst was dried at 168° C. (335°F.) in 300 cc/min nitrogen gas up-flow for 2 hours. The reactivatedcatalyst was used in a second transesterification cycle as follow.

2nd Cycle Transesterification: The transesterification was performed inidentical manner to 1st cycle. The result is illustrated in FIG. 5. Thereactivated catalyst did not perform as good as in the 1st cycle, andthe catalyst died only after about 40 hours on stream. A 2nd catalystreactivation was then performed as follow.

2nd Catalyst reactivation: After draining material from the reactor, thecatalyst in the reactor was washed with dried toluene at 10 ml/minup-flow at ambient temperature for 30 min and then the toluene wasdrained from the reactor. The catalyst in the reactor was dried at 168°C. (335° F.) in nitrogen gas at 250 cc/min up-flow for 1 hour. Asolution prepared by mixing 8 ml water, 500 ml ethanol and 1100 mltoluene was passed though the reactor at 2.2 ml/min up-flow at 168° C.(335° F.) and 12 bar (160 psig) for 12.1 hours. After draining excesssolution from the reactor, the catalyst was dried at 168° C. (335° F.)under ambient pressure in 200 cc/min nitrogen up-flow for 1 hour. Atitanium n-butoxide solution (135 g titanium n-butoxide in 2 literstoluene) was circulated through the reactor at 15 ml/min up-flow at roomtemperature for 20 minutes and then at 168° C. (335° F.) and 10.7 bar(140 psig) for 6 hours. After cooling, excess titanium solution wasdrained from the reactor. The catalyst was washed with 4 ml/min tolueneup-flow for 1.5 hours. The washed catalyst was dried at 168° C. (335°F.) in 300 cc/min nitrogen gas up-flow for 2 hours. The reactivatedcatalyst was the used in a third cycle of the transesterificationreaction as follow.

3rd Cycle Transesterification: The transesterification was performed inidentical manner to 1st cycle. The result is illustrated in FIG. 5. Thereactivated catalyst performed as well as the catalyst in the 1st cycle.Nevertheless the catalyst died after about 90 hours on stream.

The deactivated catalyst in the reactor was subjected to two morecatalyst reactivation under similar condition and then two moretransesterifications, with similar results. After the 5th cycle oftransesterification reactions, the catalyst was subjected to a 5thcatalyst reactivation, as follow. The history of catalyst reactivationsfrom the 3rd to 5th is described below.

3rd Catalyst Reactivation. After draining out all the material in thereactor leftover from 3rd cycle transesterification, the catalyst in thereactor was washed with dried toluene 10 ml/min up-flow at ambienttemperature for 1 hour and then excess toluene was drained from thereactor. The catalyst in the reactor was dried at 157° C. (315° F.) innitrogen gas at 250 cc/min up-flow for 1 hour. A solution prepared bymixing 8 ml water, 500 ml ethanol and 1100 ml toluene was passed thoughthe reactor at 2.2 ml/min up-flow at 157° C. (315° F.) and 2.7 bar (25psig) for 12.1 hours. After draining excess solution from the reactor,the catalyst was dried at 149° C. (300° F.) under ambient pressure in200 cc/min nitrogen up-flow for 1 hour. A titanium n-butoxide solution(135 g titanium n-butoxide in 2 liters toluene) was circulated throughthe reactor at 15 ml/min up-flow at room temperature for 20 minutes andthen at 157° C. (315° F.) and 7.2 bar (90 psig) for 6 hours. Aftercooling, the excess solution was drained from the reactor. The catalystwas washed with 4 ml/min toluene up-flow for 1.5 hours. The washedcatalyst was dried at 163° C. (325° F.) in 300 cc/min nitrogen gasup-low for 2 hours.

The reactivated catalyst was the subject to the 4th transesterification.The performance of the reactivated catalyst was similar to the 2nd cycletransesterification. The result is not shown in FIG. 5.

4th Catalyst Reactivation: After draining material in the reactorleftover from the 4th cycle transesterification, the catalyst in thereactor was washed with dried toluene at 10 ml/min up-flow at ambienttemperature for 1 hour and then the excess toluene was drained from thereactor. The catalyst in the reactor was dried at 157° C. (315° F.) innitrogen gas at 250 cc/min up-flow for 1 hour. A solution prepared bymixing 8 ml water, 500 ml ethanol and 1100 ml toluene was passed thoughthe reactor at 2.2 ml/min up-flow at 157° C. (315° F.) and 2.7 bar (25psig) for 12.1 hours. After draining excess solution from the reactor,the catalyst was dried at 149° C. (300° F.) under ambient pressure in200 cc/min nitrogen up-flow for 1 hour. A titanium n-butoxide solution(135 g titanium n-butoxide in 2 liters toluene) was circulated throughthe reactor at 15 ml/min up-flow at room temperature for 20 minutes andthen at 157° C. (315° F.) and 7.2 bar (90 psig) for 6 hours. Aftercooling, excess solution was drained from the reactor. The catalyst waswashed with 4 ml/min toluene up-flow for 1.5 hours. The washed catalystwas dried at 163° C. (325° F.) in 300 cc/min nitrogen gas up-low for 2hours. The reactivated catalyst was then used in a 5th cycle of thetransesterification. The performance of the catalyst was similar to 3rdcycle transesterification. The result is not shown in FIG. 5.

5th Catalyst Reactivation: After draining material from the reactor, thecatalyst in the reactor was washed with 10 ml/min dried toluene up-flowat ambient temperature for 1 hour and then excess toluene was drainedfrom the reactor. The catalyst in the reactor was dried at 124° C. (255°F.) in nitrogen gas 250 cc/min up-flow for 1 hour. Water was passedthrough the reactor at 0.3 ml/min down-flow at 152-154° C. (305-310° F.)and ambient pressure for 6 hours. The steam treated catalyst in thereactor was dried with 100 cc/min nitrogen gas down-flow for 1 hour 20minutes at 146-149° C. (295-300° F.). A titanium n-butoxide solution(135 g titanium n-butoxide in 1600 ml toluene) was circulated throughthe reactor at 15 ml/min up-flow at room temperature for 20 minutes andthen at 127° C. (260° F.) and 3.1 bar (30 psig) for 6 hours. Aftercooling, excess solution was drained from the reactor. The catalyst waswashed with 4 ml/min toluene up-flow for 1.5 hours. The washed catalystwas dried at 138° C. (280° F.) in 300 cc/min nitrogen gas up-flow for 2hours. The reactivated catalyst was then subject to a 6thtransesterification as follow.

6th Cycle Transesterification: The transesterification was performed inidentical manner to 1st cycle. The result is illustrated in FIG. 5. Thereactivated catalyst performed as well as in the 1st cycle.Interestingly, the catalyst deactivated at a slower rate.

The above experiments demonstrate that it is possible to reactivatedeactivated titanium alkoxide catalyst immobilized on a silica gelsupport in situ. However, the catalyst cycle length may be too short topractice this technology in large commercial reactors for the continuousproduction of aromatic carbonates.

Experiment 4

The objective of this experiment was to demonstrate the extendedcatalyst cycle length attainable by adding a trace amount (42 ppm Ti byweight) of a soluble Ti compound (titanium n-butoxide) into the feedstream. The deactivated catalyst from 6th cycle transesterification inExperiment 3 was again subjected to a 7th catalyst reactivation asfollows. After draining material in the reactor, the catalyst in thereactor was washed with 10 ml/min dried toluene up-flow at ambienttemperature for 1 hour and then excess toluene was drained from thereactor. The catalyst in the reactor was dried at 124° C. (255° F.) in250 cc/min nitrogen gas up-flow for 1 hour. A mixed solution of water (4wt %) in ethanol was passed through the reactor at 1.4 ml/min down-flowat 154° C. (310° F.) and ambient pressure for 6 hours. The catalyst inthe reactor was dried with 150 cc/min nitrogen gas down-flow for 1 hour25 minutes at 154° C. (310° F.). A titanium n-butoxide solution (67.5 gtitanium n-butoxide in 800 ml toluene) was passed through the reactor at15 ml/min up-flow at room temperature for 20 minutes and then at 127° C.(260° F.) under 3.4 bar (35 psig) for 6 hours. After cooling, excesssolution was drained from the reactor. The catalyst was washed with 4ml/min toluene up-flow for 1.5 hours. The washed catalyst was dried at138° C. (280° F.) in 300 cc/min nitrogen gas up-flow for 2 hours. Thereactivated catalyst was subject to the 7th transesterification cycle asfollows.

The 7th transesterification cycle was carried out by adding a traceamount (42 ppm Ti by weight) of titanium n-butoxide into the feedstreams under various reaction conditions. Two different feed mixtureswere used in this experiment. A mixed feed solution of 19.73 wt % phenoland 80.27 wt % DEC (3.24 mole ratio of DEC/PhOH) was used for the first593 hours on stream, followed by 25.83 wt % phenol and 74.17 wt % DEC(2.29 mole ratio of DEC/PhOH) for the remainder of the run until shutdown at 749 hours on stream. Titanium n-butoxide was blended into thepremixed DEC/PhOH feed solutions. The feed rates were 0.2 ml/min for thefirst 353 hours on stream time, 0.3 ml/min from 353 to 401 hours onstream and then 0.2 ml/min to the end of run. The trace analyses of theproduct samples taken at various on stream time indicated 21 ppm Ti at48 hours, 44 ppm Ti at 305 hours, 44 ppm Ti at 449 hours, 31 ppm Ti at491 hours, 51 ppm Ti at 593 hours, 51 ppm Ti at 713 hours and 31 ppm Tiat 749 hours. The result of this experiment is illustrated in FIG. 6.The temperatures cited in FIG. 6 were the temperature readings at thebottom of the catalyst bed. The temperature readings at the top of thecatalyst bed were usually 1.5-3° C. (3-5° F.) lower than the reactorbottom temperature, depending on the ethanol concentration in theproduct stream, indicating vaporization of ethanol at the top section ofcatalyst bed. Lower reactor effluent temperatures at the top of thecatalyst bed became noticeable if the ethanol concentrations in theproduct were higher than about 1.2 wt %. Phenetole was the onlydetectable by-product. The phenetole selectivity based on phenol wasless than 0.3 mole %.

As shown in FIG. 6, there was no catalyst deactivation during the entirerun time (749 hours), successfully demonstrating that the catalyst cyclelength could be extended from less than 80 hours to 749 hours or more byadding soluble 42 ppm Ti by weight into the feed stream.

The catalyst in the reactor was analyzed at the end of the run, and wasmostly yellow granules, with some catalyst granules dark brown in color.Titanium phenoxide has an intense orange or amber like color forcomparison. The analysis of the spent catalyst indicated 0.55 wt % Ti onthe catalyst. This was a surprising discovery.

Experiment 5

The objectives of this experiment were to demonstrate (1) the need ofpre-forming the catalyst prior to performing reactions, (2) catalystreactivation, (3) extension of the catalyst cycle time, and (4) the needfor controlling the water content (less than about 650 ppm by weight) infeed. In this experiment, silicon oxide pellets were used to prepare thesupport onto which titanium n-butoxide was grafted.

A silicon oxide pellet support (0.3 cm (⅛ inch), 555 ppm Na and 2500 ppmAl by weight, 280 m²/g BET SA and 1 cc/g PV) was used to prepare animmobilized titanium n-butoxide catalyst. 100 g of silicon oxide pelletswas treated at about 52° C. for 5 minutes with stirring with a sodiumhydroxide solution (10 g NaOH in 570 ml water). The silica was washedwith cold water thoroughly and then with hot water (about 80° C.) toremove trace amounts of sodium on the silica. The treated silica wasfirst dried at room temperature, then dried at 130° C. for 1.5 hours andthen at 150° C. for 1 hour in a vacuum oven. The dried silica supporthad 150 ppm Na by weight. The prepared silica support had the followingproperties: 252 m²/g BET, 1.035 cm³/g pore volume and 15.7 nm averagepore diameter.

25 ml of the dried silicon oxide pellets (9.72 g) was loaded in thereactor. A reservoir for the catalyst solution was filled with atitanium n-butoxide solution prepared by dissolving 135 g titaniumn-butoxide in 1600 ml toluene. This catalyst solution was circulatedup-flow through the reactor at a flow rate of 15 ml/min at ambienttemperature for 20 minutes, and then at 135° C. (275° F.) and 3.4 bar(35 psig) for 6 hours. After cooling, the excess catalyst solution wasdrained from the reactor and then the in situ-prepared catalyst waswashed up-flow with toluene at ambient temperature at a flow rate of 4ml/min for 1.5 hours. After draining excess toluene from the reactor,the catalyst was dried at 138° C. (280° F.) for 2 hours in 350 cc/minnitrogen gas up-flow. The resulting catalyst was used in a 1st cycletransesterification as follows.

1st Cycle Transesterification: The 1st cycle transesterification wasperformed, without injecting soluble titanium species into the feedstream, under boiling reaction conditions of 168° C. (335° F.) and 2.4bar (20 psig) with a feed rate of 0.2 ml/min up-flow. The feedcomposition was 26.07 wt % phenol and 73.93 wt % DEC (2.56 mole ratio ofDEC/phenol). The result is illustrated in FIG. 7. The catalystdeactivated with on-stream time. After about 100 hours on stream, thecatalyst had little activity.

1st Catalyst Reactivation: After draining material from the reactor, thecatalyst in the reactor was washed with dried toluene 10 ml/min up-flowat ambient temperature for 1 hour and then excess toluene was drainedfrom the reactor. The catalyst in the reactor was dried at 124° C. (255°F.) in nitrogen gas flowing at 250 cc/min up-flow for 1 hour. A mixedsolution of water (4 wt %) and ethanol was passed though the reactor at2.2 ml/min up-flow at 154° C. (310° F.) and ambient pressure for 6hours. The catalyst was dried at 154° C. (310° F.) in 150 cc/minnitrogen up-flow for 1 hour 25 minutes. A titanium n-butoxide solution(67.5 g titanium n-butoxide in 800 ml toluene) was circulated throughthe reactor at 15 ml/min up-flow at room temperature for 20 minutes, andthen at 134° C. (275° F.) and 3.4 bar (35 psig) for 6 hours. Aftercooling, the excess solution was drained from the reactor. The catalystwas washed with 4 ml/min toluene up-flow for 1.5 hours. The washedcatalyst was dried at 138° C. (280° F.) in 300 cc/min nitrogen gasup-flow for 2 hours and used in a second cycle of thetransesterification reaction.

2nd Cycle Transesterification: The reactivated catalyst was subject tothe 2nd cycle transesterification with the same feed solution atidentical conditions to the 1st cycle transesterification. The result isillustrated in FIG. 7. Similar results to the 1st cycletransesterification were obtained.

2nd Catalyst Reactivation. The 2nd catalyst reactivation was performedin identical manner to the 1st catalyst reactivation.

3rd Cycle Transesterification. The reactivated catalyst obtained fromthe 2^(nd) catalyst reactivation was subject to a 3rd cycle of thetransesterification with soluble titanium species added to the same feedsolution at identical conditions to the 1st cycle transesterification.The results of the 3^(rd) cycle transesterification are illustrated inFIG. 7. A similar result to the 1st cycle transesterification wasobtained, but the catalyst maintained a steady catalyst activity for anextended period of time. Trace analysis of a sample taken 270 hours onstream indicated 47 ppm Ti by weight. After the sample was taken at 270hours on stream, the feed reservoir was refilled with a new feed.Unfortunately, the feed became turbid when it was blended with titaniumn-butoxide. It is believed that the turbidity was caused by unexpectedlyhigher water content in feed solution than the previous feed solutions.The catalyst activity declined fast with this new feed solution. Thetrace analysis of the composite product with this new feed indicated 9ppm Ti by weight. It was discovered that the water content in the feedstream should be kept at less than about 650 ppm by weight.

Blank Run (without performing immobilization of Ti alkoxide catalyst):The same reactor was loaded with 25 ml (9.54 g) silicon oxide pelletsupports prepared by treating silicon oxide pellets with a sodiumhydroxide solution (10 g NaOH in 570 ml water) at about 52° C. for 5minutes with stirring. The silica was washed with cold water thoroughlyand then with hot water (about 80° C.) to remove trace amounts of sodiumon the silica. The treated silica was first dried at room temperature,then dried at 130° C. for 1.5 hours and then at 150° C. for 1 hour in avacuum oven. A catalyst was not grafted to the support. Thetransesterification reaction was performed with 42 ppm Ti by weight inthe same composition feed as above under identical conditions (underboiling reaction conditions of 168° C. (335° F.) and 2.4 bar (20 psig)with a feed rate of 0.2 ml/min up-flow, where the feed composition was26.07 wt % phenol and 73.93 wt % DEC (2.56 mole ratio of DEC/phenol).The results are illustrated in FIG. 7. The conversion of phenol was lessthan 2% throughout the run.

This series of experiment (Experiment 5) successfully demonstrates thatit is possible to reactivate deactivated catalyst and prolong thecatalyst cycle time to more than 250 hours. The blank run clearlydemonstrates that it is necessary to prepare the catalyst prior toperforming transesterification. Alternatively, one may choose toinitiate transesterification with a pre-prepared grafted titaniumalkoxide catalyst outside the reactor. This experiment also indicatesthat it may be necessary to control water content in the feed to lessthan about 650 ppm by weight to maintain a steady catalyst activity.

Experiment 6

The objective of this experiment was to demonstrate continuousproduction of aromatic carbonates in a series of multiple reactors inthe presence of a titanium oxide catalyst supported on a silica support.The same granular silica gel (40.7 g) in Experiment 3 was treated with asodium hydroxide solution (6.86 g NaOH in 500 ml water) at ambienttemperature for 7 minutes, with stirring. The silica gel was firstwashed with cold water thoroughly, and then with hot water (about 80°C.), to remove trace amounts of sodium on the silica. The treated silicagel was dried at 140° C. for 2 hours, at 345° C. for 3 hours, and thenat 375° C. for 2 hors. 30 ml (10.99 g) was impregnated with titaniumn-butoxide solution prepared by dissolving 4.71 g titanium n-butoxide in80 ml dry toluene. The impregnated silica gel support was calcined at500° C. for 3 hours. The titanium content on titanium oxide catalystsupported on silica was 5.48% Ti by weight based on the amount oftitanium n-butoxide used. 25 ml (9.8 g) of titanium oxide catalystsupported on silica was loaded in the reactor. Transesterification ofDEC with phenol was performed under various conditions. The feeds from 0hour to 308 hours on stream time are two different mixtures of DEC andphenol. These feeds were used to perform the 1st transesterification ofDEC with phenol. The titanium content in the DEC/PhOH feed solutions(from 0 to 308 hours on stream) was 59 ppm Ti by weight, prepared byblending a stock solution of Ti(Oet)_(4−x)(OPh)_(x) (where x=˜2). Thestock solution of Ti(Oet)_(4−x)(OPh)_(x) was prepared by distillingethanol from a solution prepared by mixing an appropriate amount oftitanium tetraethoxide in a mixed solution of DEC and phenol (PhOH) (25%wt) at from 120° C. to 125° C. for about 3 hours. The feeds from 308hours to 986 hours on stream were prepared by distilling ethanol fromthe composite products of the 1st transesterification. These feeds wereused to perform 2nd transesterifications, which are equivalent to thereactions in the second reactor in series or in some stages below thefeed point of a multi-stage catalytic distillation column. The feedsfrom 986 hours to 1136 hours on stream were prepared by distillingethanol from the composite products from 2nd transesterification. Thesefeeds were used to perform 3rd transesterifications. No soluble titaniumcatalyst component was blended into the feeds for the 2nd or 3rdtransesterification. The feed compositions are listed in Table 4. Thetransesterification was performed at 185° C. (365° F.), 2.9 bar (27psig) and a feed rate of 0.24 ml/min. The result of this experiment isillustrated in FIG. 8. The conversion of phenol in FIG. 8 is the overallphenol conversion from the 1st transesterification through the 3rdtransesterification. There was no indication of catalyst deactivationthroughout the run (1362 hours continuous operation). The examination ofthe catalyst recovered from the reactor at the end of the run indicatedlittle deposition of heavy polymers. The analysis of the catalystindicated 2.3% Ti by weight, indicating about 58% loss of Ti due toleaching into the product stream. The trace analyses of Ti in theproduct streams taken at 686, 887 and 1293 hours on stream indicated 75,57 and 78 ppm Ti by weight, respectively.

The result of this experiment clearly demonstrated the continuousproduction of aromatic carbonates by using a series of multiple reactorswith long catalyst cycle time by adding a trace amount of a soluble Ticompound in the feed stream. Also, this experiment may suggest thatlarge amounts of titanium oxide on a catalyst may not be necessary, asany excess amount of titanium oxide may be washed off by forming solubleorgano titanium compounds. The catalyst cycle time is more than longenough for the time required for the catalyst reactivation. The combinedselectivity of EPC and DPC was from about 98 mole % to about 93%, basedon converted phenol and depending on the run conditions.

TABLE 4 Time on DEC/PhOH Stream mole ratio Product Stream wt. % (h) infeed EPC wt. % in feed EPC DPC  0-86 2.2 0 5.5 0.14  86-308 2.4 05.7-6.2 0.19-0.20 308-746 1.71 5.86 (1^(st) 9.2-9.7 0.44-0.48Transesterification) 746-986 1.89 5.51 (1^(st) 8.8-9.4 0.44-0.47Transesterification)  986-1173 1.534 10.28 (2^(nd)   12-12.6 0.38-0.4 Transesterification) 1173-1362 1.69 10.3 (2^(nd) 12.2-12.8 0.74-0.78Transesterification) 185° C. (365° F.), 2.9 bar (27 psig), 0.24 ml/minfeed rate and 25 ml catalyst (9.8 g)

Experiment 7

The objective of this experiment was to demonstrate continuousproduction of aromatic carbonates in a series of multiple reactors byperforming transesterification of DEC with phenol in the presence of atitanium ethoxide catalyst immobilized on a silica gel support.

This experiment consists of two parts; Experiment 7A and 7B. InExperiment 7A, immobilization of Ti ethoxide on a silica gel support wasperformed prior to performing transesterification. The feed containedvarious amounts of soluble Ti(Oet)_(4−x)(OPh)_(x) (where x=˜2) compound.In Experiment 7B, the 25 ml space in the reactor was loaded with 25 mlthe silica gel support, and the transesterification was performedwithout grafting Ti tetra-ethoxide on the silica support.

Experiment 7A

The support used for in situ catalyst preparation was spherically shapedsilica gel spheres (1.7-4 mm diameter). This silica gel support hadabout 6 hydroxyl groups per nm², 392 m²/g BET, 0.633 cm³/g pore volume,6.48 nm average pore diameter, and about 0.58 g/ml apparent bulk density(ABD). This silica gel support (25 ml; 14 46 g) was loaded in thereactor. A titanium ethoxide solution (45.25 g titanium ethoxide in 800ml toluene) was circulated up-flow at 15 ml/min through the reactor atambient temperature for 20 minutes and then at 135° C. (275° F.) under3.4 bar (35 psig) for 6 hours to graft titanium ethoxide on silica gelsupport. After cooling, excess solution in the system was drained andthen the catalyst was washed with toluene at 4 ml/min for 1.5 hours. Thewashed catalyst was dried at 138° C. (280° F.) for 2 hours in 300 cc/minnitrogen gas flow rate.

Reactions producing EPC and DPC were performed at various conditions.The result is illustrated in FIGS. 9A and 9B. All the firsttransesterification reactions were performed with 59 ppm Ti asTi(Oet)_(4−x)(OPh)_(x) (where x=˜2) added into feed stream, except theperiod of from 709 hours to 799 hours on stream, where 151 ppm Ti wasadded. The run started at 185° C. (365° F.), 2.9 bar (27 psig) and 0.24ml/min for the first transesterification. After the first 50 hrs onstream, the temperature was slowly lowered to 174° C. (345° F.) and thefeed rate was slowly increased to 0.5 ml/min over the next 96 hrs onstream. Thereafter, all the 1st and 2nd transesterification wereperformed at the condition of 174° C. (345° F.), 2.9 bar (27 psig) and0.5 ml/min. By distilling ethanol from the composite products from the1st, 2nd, and 3rd transesterifications, the feed mixtures were preparedfor the 2nd, 3rd and 4th transesterifications. From 973 hrs to 1064 hrson stream time, a mixed transesterification and disproportionationreaction was performed at 174° C. (345° F.) and 2.4 bar (20 psig), 0.5ml/min feed rate. The composition of the feed for the reaction was18.553% DEC, 0.108% ethylbutyl carbonate, 0.283% phenetole, 0182%unknown, 57.508% phenol, 22.03% EPC, 0.054% p-phenoxyphenyl methylcarbonate, and 1.282% DPC, on a weight basis. The result indicates thatthe major reaction is disproportionation. But the analysis of the dataalso suggests the need for removing DEC from the feed fordisproportionation. In FIG. 9A, the 2nd transesterifications wereperformed at 174° C. (345° F.), 2.4 bar (20 psig) and 0.5 ml/min withthe Ti concentration in the feed ranging from 44 ppm Ti to 69 ppm Ti byweight. In FIG. 9B, the 2nd transesterifications were performed at 174°C. (345° F.), 2.9 bar (27 psig) and 0.5 ml/min with the Ti concentrationin feed ranging from 45 ppm Ti to 75 ppm Ti by weight. The 3rdtransesterification was performed at 174° C. (345° F.), 2.5 bar (22psig) and 0.5 ml/min with the Ti concentration in the feed ranging from52 ppm Ti to 74 ppm Ti by weight. The 4th transesterification wasperformed at 174° C. (345° F.), 2.4 bar (20 psig) and 0.5 ml/min withthe Ti concentration in the feed ranging from 51 ppm Ti to 73 ppm Ti byweight.

The selectivity of aromatic carbonates decreased with the conversion ofphenol. The combined selectivity of EPC and DPC during the 1sttransesterification is about 99 mole % based on phenol. The combinedselectivity of EPC and DPC during the 4th transesterification was 94mole to 95 mole % based on phenol.

The solid catalyst had been running over 14 months until termination ofthe experiment unrelated to catalyst activity. FIGS. 9A and 9B stronglysuggest that there has been little to no catalyst deactivation for over14 months. The analyses of the two catalyst samples taken carefully fromthe top and bottom of the catalyst bed indicate a same amount of 0.28 wt% Ti (550° C. calcined based) on the both catalyst samples. Thisexperiment successfully demonstrates that a long catalyst cycle (morethan 14 months) can be obtained by adding a trace amount of solubletitanium compound into the feed stream.

Experiment 7B Blank Run

The objective of this experiment was an attempt to immobilize Tialkoxide on a silica support while performing the transesterification.Various amounts of soluble Ti(OEt)_(x)(OPh)_(4−x) compound were addedinto the feed stream. The transesterification of DEC with phenol wasperformed at 174° C. (345° F.) and 2.9 bar (27 psig). The result isillustrated in FIG. 10.

Comparing the Blank in Experiment 5 (FIG. 7) and Experiment 7B (FIG. 10)with the result (FIGS. 9A and 9B), there is clear need for immobilizingtitanium alkoxide on silica gel support prior to performingtransesterifications. Comparing Comparative Experiments 1 and 2 (FIG. 4)with FIGS. 9A and 9B, the superiority of the novel solid catalysttechnology, adding a trace amount of soluble active organometalliccompound to the feed, over the prior art is also clearly demonstrated.

Experiment 8

The objective of this experiment was to demonstrate disproportionationof EPC to DPC and DEC in the absence of solid catalyst, but in thepresence of soluble Ti catalytic components. The feeds for thedisproportionation were prepared by distilling ethanol, DEC and a partof phenol from the composite products from the 4th transesterificationin Experiment 7 under a nitrogen blanket. The homogeneous Ti catalyst inthe feed mixtures was originated from the 4th transesterificationcomposite products. No additional soluble Ti catalyst was added to thefeed mixtures. Toluene was added to two feed mixtures to create a vaporphase for the boiling point reactor. The first feed composition was16.26% toluene, 1.61% DEC, 49.33% phenol, 30.91% EPC and 0.78% DPC byweight and the balance was by-products including trace amounts of MPC.The second feed composition was 16.15% toluene, 1.61% DEC, 49.28%phenol, 31.08% EPC and 0.80% DPC by weight and the balance wasby-products including trace amount of MPC. The concentration ofhomogeneous catalyst in the first and second feeds was 180 ppm and 200ppm Ti by weight, respectively.

The disproportionation was performed in the reactor with 25 ml emptycatalyst space (in the absence of solid catalyst) at 179° C. (355° F.)and 2.9 bar (27 psig). The feed rates were 0.5 ml/min up-flow for thefirst 72 hours on stream and then 0.60 ml/min up-flow thereafter. Theresults of the disproportionation reactions are illustrated in FIG. 11.The experimental results indicate that small amounts of EPC are alsoproduced in addition to DPC. Xanthone was the only new by-productsproduced during disproportionation in an amount of about 35 ppm byweight. Diphenyl ether was not detected in any sample analysis. Theselectivity of all the by-products was from 3.0 mole % to 3.3 mole %.This experiment demonstrates successfully the EPC disproportionation toproduce DPC and DEC according to embodiments disclosed herein.

Experiment 9

This experiment demonstrates the purification of DPC. A compositedisproportionation product from Experiment 8 was distilled to removeethanol, DEC and a substantial portion of phenol by using laboratorydistillation equipments. The remaining material in the distillationflask had the following composition: 0.024% EtOH, 0.204% DEC, 0.017%phenetole, 1.619% unknowns, 12.563% phenol, 25.377% EPC, 59.474% DPC and0.723% heavies. By performing vacuum distillation, crude DPC (cut atvapor temperature from 235 to 245° C.) was obtained. The composition ofthis crude DPC was 0.535% unknowns, 2.112% phenol, 0.013% phenyl ether,0.030% EPC, 94.555% DPC, 0.026% xanthone and 2.73% heavies. This crudeDPC was re-crystallized in a mixture of 5 wt % diethyl ether in hexanefive times. The final DPC product had impurities of 0.4 ppm xanthone and11.6 ppm phenol by weight. No other impurities were detected by traceanalysis. This DPC product has a greater purity than high purity DPCavailable in the market (28.7 ppm unknowns and 67.2 ppm phenol byweight).

Dialkyl Carbonates by Transesterification of a Cyclic Carbonate with anAlcohol

Dialkyl carbonates are continuously produced by performingtransesterification of a cyclic carbonate with alcohols in the presenceof solid catalysts. As described above, embodiments disclosed herein maybe particularly useful for continuous production of dialkyl carbonates,such as DMC, DEC, etc. There are a number of homogeneous catalysts fortransesterification. When dialkyl carbonates are produced by performingtransesterification of a cyclic carbonate with an alcohol in thepresence of supported metal oxide or mixed metal oxide catalysts or asolid catalyst prepared by immobilizing a homogeneous catalyst on aporous support, the catalysts have unacceptably short cycle length foroperation of large commercial reactors. The permanent catalystdeactivations involved in dealing with organic carbonates are caused byleaching active catalytic components out of heterogeneous catalysts intothe reaction medium. Therefore, dialkyl carbonate such as DMC iscommonly produced by performing transesterification in the presence of ahomogeneous catalyst.

Embodiments disclosed herein, however, offer processes for producingdialkyl carbonate in the presence of a solid catalyst. Solid catalystsmay include one or more elements from Groups II, III, IV, V and VI ofthe Periodic Table. A first type of solid catalyst includes one or moreorganometallic compounds of the above elements immobilized on a poroussupport, which may have surface functional groups such as hydroxyl,carbonyl, alkoxy, a mixture of hydroxyl and alkoxy, chlorine, etc.Supports may include silica, titanium oxide, zeolitic materials such asMCM-41, MCM-48, SBA-15, carbon and/or carbonaceous materials, etc. Asecond type of solid catalyst includes metal oxides, hydroxides oroxyhydroxides of one or more of the above elements deposited on a poroussupport. To maintain a stable catalyst activity, a trace amount of asoluble catalytic component is added to the feed stream. By doing so,the catalyst cycle length can be extended to be suitable for commercialreactor.

The transesterification may be performed in any physical devices, suchas traditional fixed bed reactors, catalytic distillation columns,boiling point reactors, divided wall distillation columns, pulsed flowreactors, or any combination of these devices. Examples of a combinationmay include a fixed bed boiling point reactor followed by a catalyticdistillation column reactor. Transesterification of cyclic carbonateswith a primary alcohol, such as ethanol or methanol, may be carried outas a two-step reaction, where there are two ethyl propylene glycolcarbonate intermediates. Also the reaction product contains smallamounts of by-products, such as propylene glycol ethyl ether isomers,produced by O-alkylation of propylene glycol by DEC. Thetransesterification may be performed in a single reaction zone ormultiple reaction zone.

FIG. 18 illustrates a simplified process flow diagram for the continuousproduction of DEC and propylene glycol co-product by performingtransesterification of propylene carbonate with ethanol in the presenceof a solid catalyst according to embodiments disclosed herein. Thetransesterification, as illustrated, may be performed in a catalyticdistillation reactor 101 at a temperature from about 149° C. to about177° C. (about 300° F. to 350° F.) under a pressure from aboutsubatmospheric pressures to about 11.4 bar (i.e., about 7 psia to 165psia) depending on the composition of reaction mixture. In addition tocatalytic distillation reactor 101, the process includes twodistillation columns, 102 and 114. Catalytic distillation column 101includes a reaction zone RZ, in which a solid catalyst may be located.Fresh propylene carbonate feed 105 is combined with recycle stream 117and the combined stream 106 is introduced into the catalyticdistillation column 101 at a suitable position above the solid catalystbed reaction zone RZ.

Column 101 overheads stream 107, a mixture of ethanol, DEC and lights,such as carbon dioxide, is introduced to DEC recovery column 102, forthe separation of DEC from lighter components. Column 102 overheadsstream 108 may be introduced into gas-liquid separation drum 110 toseparate liquid ethanol from gases vented through line 111. The liquidethanol recovered from drum 110 in stream 112 is combined with the freshethanol feed stream 103 and the combined stream 104 is heated to produceethanol vapor introduced to catalyst distillation column 101 at suitableposition below reaction zone RZ. Distillation column 102 bottoms stream109 contains the product, DEC, which may be sent to a storage tank (notshown) or other downstream processes.

Bottom stream 113 from the catalytic distillation column 101, whichcontains propylene glycol, propylene carbonate, reaction intermediatesand by-products such as 1-ethoxy-2-propanol, heavies, etc. and a traceamount of catalyst, are introduced to the second distillation column 114to recover an overhead stream 115 containing propylene glycol,1-ethoxy-2-propanol, etc. Propylene glycol may be recovered from themixture in stream 115 by distillation (not shown). Column 114 bottomsstream 116 is recycled to the catalytic distillation column 101 throughlines 117 and 106. A portion of bottoms steam 116 may be purged from thesystem via stream 118 to prevent build-up of heavies in the system.

A trace amount of soluble organometallic compound is introduced to thecatalytic distillation column 101 above catalytic reaction zone throughline 119. In some embodiments, the catalyst solution is fed at a ratesuch that the liquid reaction mixture flowing down the catalyticreaction zone RZ contains a trace amount, typically from 5 ppm to about100 ppm metal by weight, of soluble metallic component, such as Mg, Ca,Zn, La, Ac, or Ti compounds.

The production of dialkyl carbonates is illustrated by the followingexperiments.

Experiment 10

The objective of this experiment was to demonstrate thetransesterification of propylene carbonate with ethanol to produce DECand propylene glycol in the presence of a solid catalyst. The solidcatalyst is prepared in situ by immobilizing titanium ethoxide on asilica gel support.

The reactor was loaded with 25 ml (1.7-4 mm diameter) of a sphericalsilica gel support. The weight of the support was 10.097 g. This silicagel support had about 6 hydroxyl groups per nm², 314 m²/g BET, 1.055cm³/g pore volume and 13.46 nm average pore diameter. A titaniumethoxide solution (40 g titanium ethoxide in 800 ml toluene) wascirculated up-flow through the reactor at 15 ml/min at ambienttemperature for 30 minutes and then at 135° C. (275° F.) and 3.4 bar (35psig) for 6 hours to graft titanium ethoxide on the silica gel support.After cooling, excess solution was drained from the reactor and then thecatalyst was washed with toluene at 4 ml/min for 1.5 hours. The washedcatalyst was dried at 138° C. (280° F.) for 2 hours in 300 cc/minnitrogen gas flow.

Mixed solutions of propylene carbonate and ethanol were prepared and 45ppm Ti as titanium ethoxide was blended into the mixed feed solutions.Transesterifications were performed with various feed mixtures inup-flow liquid phase at 174° C. (345° F.) and 17.9 bar (245 psig). Therun conditions are listed in Table 5. The results of this experiment areillustrated in FIG. 12.

TABLE 5 EtOH/PC Time on mole Feed Rate Ti in Feed Stream (h) ratio infeed (ml/min) (wt. ppm) Note  0-24 0.64 0.3 45 1^(st)Transesterification 24-90 6.64 0.4 45 1^(st) Transesterification  90-1205.23 0.4 45 1^(st) Transesterification 120-143 4.89 0.4 45 1^(st)Transesterification 143-167 3.67 0.4 45 1^(st) Transesterification167-215 6.41 0.4 45 1^(st) Transesterification 215-287 5.52 0.4 451^(st) Transesterification 287-335 5.52 0.5 45 1^(st)Transesterification 335-383 6.76 0.5 45 1^(st) Transesterification383-647 6.9 0.5 89 2^(nd) Transesterification 647-815 6.3 0.5 90 3^(rd)Transesterification 815-887 6.41 0.5 45 1^(st) Transesterification887-893 4.95 0.5 45 1^(st) Transesterification  893-1055 6.03 0.5 451^(st) Transesterification 1055-1151 6.2 0.5 45 1^(st)Transesterification 1151-1223 4.13 0.5 45 1^(st) Transesterification1223-1313 4.72 0.5 45 1^(st) Transesterification 1313-1397 3.65 0.5 451^(st) Transesterification 1397-1469 5.27 0.5 0 1^(st)Transesterification

The results of this experiment clearly demonstrate that DEC (a dialkylcarbonate) may be produced by performing transesterification of a cycliccarbonate such as propylene carbonate with ethanol in the presence ofsolid Ti alkoxide catalyst immobilized on a silica gel support by addinga trace amount of a soluble Ti compound into the feed stream. Withoutaddition of a trace amount of Ti in feed stream, the catalyst activitydeclines rapidly as shown in FIG. 12 over run hours 1397-1469.

Experiment 11

The objective of this experiment is to demonstrate transesterificationof propylene carbonate with ethanol to produce DEC and propylene glycolin the presence of a solid catalyst. The experiment consists of twoparts; Comparative Experiments 11A (non-invention) and 11B.

Comparative Experiment 11A

The transesterification was performed in the presence of a homogenousmagnesium tert-butoxide. The mole ratio of ethanol/propylene carbonateof the feed mixture was 6.85. The concentration of homogeneous catalystwas 57 ppm Mg by weight. The transesterification was performed at 168°C. (335° F.), 17.9 bar (245 psig) and 0.5 ml/min. The result isillustrated in FIG. 13. The average conversion of propylene carbonate isabout 24.3 mole %. The average selectivity of DEC and propylene glycolare 95.7 and 94.3 mole %, respectively.

Experiment 11B

The transesterification was performed in the presence of a solidcatalyst. The starting solid catalyst was MgO supported on a silica gelsupport. A magnesium nitrate solution was prepared by dissolving 10.098g of Mg(NO₃)₂.6H₂O in 22.73 g deionized water by incipient impregnation.Magnesium nitrate was deposited on 30 ml (11.777 g) of the same silicagel support used in Experiment 10 by incipient impregnation. Theimpregnation product was dried at 100° C. in vacuum oven for one hour,followed by calcinations at 510° C. for 2 hours to prepare MgO supportedon silica gel. 25 ml (10.77 g) of this surface mixed oxide catalyst ofMgO and silica gel was loaded in the reactor. Transesterification ofpropylene carbonate with ethanol was performed at the various conditionslisted in Table 6.

TABLE 6 EtOH/PC Mg in Time on mole Feed Rate Feed Stream (h) ratio infeed (ml/min) (wt. ppm) Note  0-79 6.96 0.5 111 1^(st)Transesterification  79-175 10.12 0.5 111 1^(st) Transesterification175-341 6.95 0.5 56 1^(st) Transesterification 341-413 6.89 0.5 34.81^(st) Transesterification 413-485 7.17 0.5 42.8 1^(st)Transesterification 485-570 3.41 0.5 41.5 1^(st) Transesterification570-666 4.14 0.5 34.5 1^(st) Transesterification 666-738 5.39 0.5 34.51^(st) Transesterification 738-834 6.59 0.5 37.5 2^(nd)Transesterification

The reaction products contained 1-ethoxy-2-propanol and dipropyleneglycol as by-product. No diethyl ether was detected in any productsample. The result is also illustrated in FIG. 14. The averageselectivity of DEC and propylene glycol for the 1st transesterificationare 95.3 mole % and 94.8 mole %, respectively. In general, theselectivity declines slowly with conversion of propylene carbonate. Alsoselectivity increases with the mole ratio of EtOH/propylene carbonate.The average selectivity of DEC and propylene glycol for the 2ndtransesterification are 94.0 mole % and 92.8 mole %, respectively.

Dialkyl Carbonates from Urea and an Alcohol According to EmbodimentsDisclosed Herein

According to publications such as P. Ball et al. and D. Wang et al., asreferenced above, heterogeneous catalysts useful for the production ofdialkyl carbonates from urea and an alcohol may include Al₂O₃, Sb₂O₃,and silica. Fused SiO₂ is not a catalyst, but may become catalytic inthe presence of PPh₃. ZnO and MgO supported on silica may also be usedto produce dialkyl carbonates from urea and an alcohol.

However, metal oxide catalysts such as ZnO or MgO leach out of solidcatalysts under reaction condition, resulting in permanent catalystdeactivation. Catalyst cycle length is very important in commercialproduction of dialkyl carbonates using a catalytic distillation columnreactor, as catalytic distillation provides for fast removal of DMC orDEC and ammonia from the liquid catalytic reaction medium, improving thedialkyl carbonate productivity and selectivity. Additionally, the abovedescribed heterogeneous catalysts are not as effective as homogeneousdibutyltin dimethoxide catalyst.

According to embodiments disclosed herein, dialkyl carbonates may becontinuously produced by causing the alcoholysis of urea with an alcoholin two steps in the presence of a solid catalyst. Both reaction stepsare equilibrium reactions. Examples of alcohols used to produce thedialkyl carbonates include methanol, ethanol, propanol, etc. In thefirst step of the reaction, urea is reacted with an alcohol in areactive distillation column (pre-reactor), serving as the firstreaction zone, in either the absence or presence of a catalyst toproduce alkyl carbamate and ammonia. Catalyst is not necessary for thefirst step reaction. Impurities in feed streams such as water andammonium carbamate are removed as CO₂ and ammonia in the first reactionzone as well, protecting downstream catalysts. In the second stepreaction, alkyl carbamate produced in the first reaction zone is reactedwith alcohol to produce dialkyl carbonate and ammonia in the presence ofa solid catalyst in one or more catalytic distillation columns (primaryreactors), which are serving as the second reaction zone. The two stepreaction may be illustrated as follows:

The solid catalyst may be prepared by immobilizing, for example, anorganotitanium compound on a support such as silica or carbonaceousmaterial. There are two types of catalysts at the beginning of thereaction. The first type of catalyst is metal alkoxide, a metal salt ofa carbonic acid monoester, or a mixture of these, immobilized on asupport such as silica or carbonaceous material. The second type of thecatalyst is metal oxide supported on a support such as silica, alumina,titania, zirconia, carbonaceous material, etc. Active metal componentsmay be elements such as Sn, Sb, Ti, Zr, Zn, Mg, Ca, etc.

Again, a small amount of a soluble metal compound is added to thereaction stream going into the reactor to maintain catalyst activityover an extended cycle length. By doing so, the catalyst cycle lengthcan be extended to be suitable for use in commercial processes. It isbelieved that the working catalyst under the steady state condition is ametal alkoxide, metal alkoxy alkyl carbonate (salt of carbonic acidmonoester, or oligomers or mixtures of these, immobilized on a support.The concentration of soluble organometallic compound, such as dibutyltindialkoxide, in the reaction mixture is significantly lower than thehomogeneous catalyst used in U.S. Pat. No. 7,074,951.

A high boiling solvent, such as triglyme, may be used as a solvent inthe second step, and serves as a co-catalyst to improve reaction rateand selectivity. Importantly, because of the high boiling point of thesolvent, the reaction can be carried out under low pressure, which aidsin fast removal of DEC and ammonia from the liquid reaction medium intothe vapor phase along with excess ethanol vapor as stripping gas,resulting in high DEC productivity and selectivity. Embodimentsdisclosed herein offer an alternative improved process for producingdialkyl carbonate in the presence of a solid catalyst. This process maybe a “greener” process, because the concentration of tin catalyst in theprocess stream is substantially reduced compared with homogeneouscatalyst-based processes as demonstrated in Example 14 below. The traceamount of soluble catalytic compound in the process stream stays in thesystem. No recovery or separation of soluble catalyst compound in theprocess stream is necessary.

FIG. 15 illustrates a flow diagram of the process for the continuousproduction of DEC according to embodiments disclosed herein. A doublediameter distillation column reactor 36 is used as the pre-reactor toremove impurities in the feed streams and for the conversion of urea toethyl carbamate (EC). Urea solution is prepared in drum 34 by mixingurea feed 31 and ethanol stream 33. Ethanol stream 33 may include freshethanol feed 32 and ethanol from recycle stream 74.

The urea solution 35 from drum 34 is introduced to the middle of anupper narrower column section of the double diameter tower reactor 36.Reactor 36 serves as a pre-reactor to clean up the impurities (water andammonium carbamate) in the feeds, ethanol and urea, and to convert ureato EC. Vapor stream 37 from pre-reactor 36 is composed of ammonia,carbon dioxide and ethanol. The cleaned mixed solution is removed frompre-reactor 36 as bottom stream 38. Stream 38 is introduced to primaryreactor 39 (catalytic distillation column) at a position above acatalytic reaction zone 39R containing a solid catalyst.

Recycle ethanol stream 40 is introduced into reactor 39 as superheatedethanol vapor at a position below the catalytic reaction zone 39R. Thebottom stream from the catalytic distillation column 39 is recycled to aposition above the feed point of line 38 at the top of column 39 throughlines 42, 44 and 78. A small slipstream 43 from the recycle loop iscombined with bottom stream 65 from DEC recovery column 63 to stream 66,which is introduced to clean-up reactor 67 at a position above catalyticreaction zone, which is another small catalytic distillation columncontaining a solid catalyst. Slipstream 43 may include ethanol, ammonia,ethyl amine, diethyl ether, DEC, ethyl carbamate, N-ethyl ethylcarbamate, triglyme (TG), heavies and a trace amount of soluble catalystcomponent. The bottom stream 65 from DEC recovery column 63 may includeethyl carbamate, N-ethyl ethyl carbamate, TG, and a trace amount ofcatalyst. The overhead stream 68 from clean-up reactor 67 may includeammonia, ethyl amine, CO₂, diethyl ether, ethanol, and DEC. Bottomstream 69 from clan-up reactor 67 may include ammonia, ethyl amine, CO₂,diethyl ether, ethanol, N-ethyl ethyl carbamate, ethyl carbamate,heterocyclic compounds and a trace amount of soluble catalyst component.

Bottom stream 69 from reactor 67 is cooled to precipitate heterocycliccompounds in the cooling/filter system 70. The precipitated solidby-product is removed from system 70 through line 71. Liquid stream 72from system 70 splits to two streams 77 and 78 to recycle to clean-upreactor 67 and primary reactor 39, respectively.

Overhead stream 41 from primary reactor 39 may be combined with overheadstream 68 from clean-up reactor 67 to stream 42. Overhead stream 41 fromprimary reactor 39 may include ammonia, CO₂, ethyl amine, diethyl ether,ethanol, ethyl carbamate, N-ethyl ethyl carbamate, DEC, TG and traceamount of catalyst. Combined stream 42 is introduced to distillationcolumn 43, where lights and heavier compounds are separated. Overheadstream 44 from distillation column 43, which may include ammonia, CO₂,ethylamine, diethyl ether and ethanol, is combined with overhead stream37 from pre-reactor 36 to stream 45 to introduce to distillation column46.

Overhead stream 47 from distillation column 46 is cooled to cause thereaction of CO₂ with ammonia to form ammonium carbamate. Ammoniumcarbamate is precipitated in liquid ammonia and removed as solidsthrough line 49 from cooling/filter system 48. Liquid ammonia stream 50from cooling/filter system 48 is sent to an ammonia storage tank.

Bottom stream 51 from column 46 may include ethylamine, diethyl ether,ethanol and a trace of DEC. Stream 51 is introduced to ethylaminerecovery column 52. Overhead ethylamine stream 53 is sent to a storagetank. Bottom stream 54 from column 52 is combined with bottom stream 55from distillation column 43 to stream 56. Stream 56 is introduced toether recovery column 57. Ether is removed from distillation column 57as overhead stream 58, which is sent to an ether storage tank. Bottomstream 59 from distillation column 57 is introduced distillation column60 (ethanol recovery column).

Recovered ethanol as overhead stream 61 is recycled to primary reactor39, clean-up reactor 67 and pre-reactor 36 (or drum 34). The ethanolrecycle stream 74 is a minor portion of the overhead stream 61 ofdistillation column 60 (ethanol recovery column). Stream 61 splits tothree streams, 40, 73 and 74. Stream 73 is recycled to clean-up reactor67. Stream 74 is recycled to drum 34 for preparation of urea solution.Stream 40, which may be a major portion of stream 61, is recycled toprimary reactor 39. Bottom stream 62 from ethanol recovery column 60 isintroduced to DEC recovery column 63. Product DEC is recovered asoverhead stream 64 from distillation column 63 and sent to a DEC storagetank. The bottom stream 65 from column 63 may include ethyl carbamate,N-ethyl ethyl carbamate, TG and a trace amount of soluble catalystcomponent. This stream 65 is sent to clean-up reactor 67 via line 66.

DMC can be produced from methanol and urea in a process similar to theprocess for producing DEC as described with respect to FIG. 15. However,it is understood that the final product DMC is recovered from a processstream having a methanol-DMC azeotrope. Recovery of DMC by breaking amethanol-DMC azeotrope by a solvent extractive distillation technique iswell documented, such as described in U.S. Pat. No. 7,074,951.

Experiment 12

The objective of this experiment is to demonstrate the reaction of ethylcarbamate with ethanol to produce DEC and ammonia in the presence of asolid catalyst. The solid catalyst was prepared by immobilizingdibutyltin dimethoxide on a silica gel by in situ technique.

25 ml (14.79 g) of spherically shaped silica gel support used in theExperiment 7A was loaded in the reactor. A dibutyltin dimethoxidesolution was prepared by mixing 87 g dibutyltin dimethoxide in 2 litersof dry toluene. The reactor was filled with this solution up-flow atambient temperature and pressure. The reactor was heated slowly to 110°C. (230° F.) while flowing this solution up-flow at 2 ml/min. At 110° C.(230° F.), the reactor was placed under 3.4 bar (35 psig) and thencontinued to heat to 135° C. (275° F.). At 135° C. (275° F.) and 3.4 bar(35 psig), the dibutyltin dimethoxide solution was passed through thereactor up-flow at 0.5 ml/min for 6 hours. After cooling, excesssolution in the reactor was drained and then the catalyst was washedwith dry toluene at 4 ml/min up-flow for 1.5 hours. The washed catalystwas dried at 104° C. (220° F.) under ambient pressure in 300 cc/min N₂flow down-flow for 2 hours.

The reaction was performed by passing a solution of 13.2% ethylcarbamate, 31.36% triglyme and 55.44% ethanol by weight up-flow withnitrogen gas through the solid catalyst bed in the boiling pointreactor. One may also carry out the reaction in down-flow reactor. Traceamounts of dibutyltin dimethoxide were blended into this solution. Thereaction conditions are listed in Table 7 and the results of this testare illustrated in FIG. 16. Analyses of the reaction products indicatedtrace amounts of N-ethyl ethyl carbamate and diethyl ether. Theselectivity of DEC based on ethyl carbamate was in a range of from 98.5mole % to 99.9 mole % with a general trend of decreasing selectivity asconversion of ethyl carbamate increases.

TABLE 7 Pressure, Time on Stream Temperature, bar Feed Rate N₂ Flow RateDibutyl Tin Dimethoxide (h) ° C. (° F.) (psig) (ml/min) (cc/min) (ppm Snby weight)  0-72 174 (345) 6.2 (75) 0.5 1.5 650 72-96 174 (345) 6.2 (75)0.5 1.5 1300  96-156 174 (345) 6.2 (75) 0.5 1.8 1300 156-254 174 (345)6.2 (75) 0.5 1.5 1950 254-298 174 (345) 5.4 (63) 0.5 0 1950 298-406 174(345) 5.4 (63) 0.5 1.0 1950 406-456 174 (345) 5.4 (63) 0.5 1.5 1950

This experiment successfully demonstrates that DEC may be produced fromurea and ethanol. In the first step, ethyl carbamate was produced byreacting urea with ethanol in the absence of a catalyst (see U.S. Pat.No. 7,074,951). In the second step, DEC was produced by performing thereaction of ethyl carbamate with ethanol in the presence of a solidcatalyst and with addition of a trace amount of soluble organometalliccompound to the feed streams to counter balance the loss of metal due toleaching. The commercial production of DEC in the second step ispreferably performed in one or more catalytic distillation columns.

Integration of Dialkyl Carbonate and Diaryl Carbonate Processes

An integrated process for producing diaryl carbonates is disclosed,where the process does not require a solvent-based extractivedistillation unit for the separation of azeotropic process streams,saving energy and construction costs, which reduces the emission of GHG(greenhouse gas) CO₂ to atmosphere. Therefore, the integrated process isa greener process compared to conventional processes for the productionof diaryl carbonates.

A greener process may be achieved by integrating a process for producingdialkyl carbonates and a process for producing diaryl carbonatesaccording to embodiments disclosed herein. For example, dialkylcarbonates may be produced in the front end of the integrated process,at least a portion of which may then be used in the back end of theintegrated process to produce diaryl carbonates. As described above,dialkyl carbonates may be produced via one of two processes: (1)transesterification of a cyclic carbonate as may be produced from anepoxide and carbon dioxide, and (2) reaction between urea and analcohol, where the urea may be synthesized using ammonia and carbondioxide. The dialkyl carbonates produced in the front end of theintegrated process via at least one of the above two processes may thenbe further reacted with an aryl hydroxy compound in the back end of theintegrated process in order to produce the diaryl carbonates.

Thus, integrated processes for the production of diaryl carbonatesaccording to embodiments disclosed herein may be produced from dialkylcarbonates exclusively produced from ethanol and CO₂. The co-productethanol from the process step of transesterification of DEC to EPC(ethyl phenyl carbonate) is recycled to produce DEC. EPC isdisproportionated to produce DPC and co-product DEC. The DEC is producedfrom carbon dioxide and ethanol via intermediates. No extractivedistillation unit is needed to produce DEC, because ethanol and DEC donot form an azeotrope (as compared to methanol and DMC, which do form anazeotrope). The processes for producing DPC from carbon dioxide andphenol may be performed using various multi-step reaction routes, whichare detailed further below.

The advantages of the integrated processes for producing DPC from DECaccording to embodiments disclosed herein may include energy saving andmaterial saving for the construction of a plant. Processes according toembodiments do not require separation of materials from compositionsthat form an azeotrope, and thus the equipment and energy usage forprocesses disclosed herein may be reduced as compared to producing DPCfrom DMC, which results in an azeotrope (DMC and methanol).

Using an integrated process for producing DEC and DPC according toembodiments disclosed herein may additionally provide several advantagesover the stand-alone processes. One such advantage of the integration isthat the ethanol co-product produced in the back-end DPC process may berecycled back as feedstock for the front-end DEC process. Therefore, theoverall material feedstock costs may be reduced. In some embodiments,all or a portion of the recycled ethanol co-product may be combined witha fresh ethanol make-up stream prior to feeding the DEC process. Inother embodiments, the recycled ethanol co-product alone may besufficient to feed the DEC process. Alternatively, one may choose to usebio-ethanol instead of synthetic ethanol to produce DEC at the start-upof an integrated plant of producing DPC and/or as make-up ethanol.

Another advantage of using the integrated DEC and DPC processes is theenergy cost and material construction costs savings associated withdrying the ethanol feedstock to the DEC process. For example, ethanol isa hydroscopic compound that tends to absorb water from the surroundings,such as atmospheric moisture. The water impurities contained in theethanol feed can adversely affect the DEC process, for example, bypoisoning or deactivating the catalyst and by plugging the processlines. Therefore, the fresh ethanol feed to the DEC unit typicallyrequires drying to produce a high-purity ethanol by separating out thewater impurities. In contrast, the ethanol co-product recovered from theback-end DPC process may contain little to no water impurities.Therefore, if the ethanol is recycled to the DEC process, the dryingcosts, including the material construction and the energy costs, may besubstantially reduced or even completely eliminated.

Yet another advantage of using an integrated DEC and DPC process may bethe construction and the operating cost savings resulting fromeliminating the redundancy in the processing equipment. For example, theDEC and DPC process may each require a separator for separating DEC fromethanol. In the DEC process, the reactor effluent may contain both DECand unreacted ethanol, which may require separation. In the DPC process,the light effluent recovered from the transesterification may containboth DEC and ethanol, which also may require separation. However,integration of DEC and DPC processes according to embodiments disclosedherein may allow use of a single system for the separation of DEC andethanol, thus resulting in construction and operating cost savings.

One method for producing diaryl carbonates according to embodimentsdisclosed herein involves 1) reacting carbon dioxide with an epoxide toproduce a cyclic carbonate, 2) transesterifying the cyclic carbonatewith ethanol to produce diethyl carbonate and 3) transesterifying thediethyl carbonate to form an ethyl aryl carbonate, and 4)disproportionating the ethyl aryl carbonate to from diaryl carbonate.Each of these reactions may be performed in one or more reaction zones,intermediate or inclusive of which may be one or more separation stagesfor separating reaction products, reactants, and/or reactionby-products.

For example, carbon dioxide available at an ammonia plant, a syngasplant, or an electric generator, may be contacted with an epoxide, suchas at least one of ethylene oxide and propylene oxide, in a cycliccarbonate synthesis reactor to produce a cyclic carbonate. An effluentfrom the cyclic carbonate synthesis reactor containing a cycliccarbonate may be contacted with ethanol in the presence of atransesterification catalyst according to embodiments disclosed hereinto produce diethyl carbonate and a glycol. The diethyl carbonate productand unreacted ethanol may be recovered and separated. The diethylcarbonate may then be fed to the transesterification reaction forproducing ethyl phenyl carbonate, and for further producing the diphenylcarbonate from the ethyl phenyl carbonate via a disproportionationreaction according to embodiments disclosed herein. The ethanol may bereturned to the system for transesterifying a cyclic carbonate toproduce diethyl carbonate.

Another method for producing diary carbonates according to embodimentsof disclosed herein involves urea, which may be produced by reactingcarbon dioxide with ammonia. Alcoholysis of the urea with ethanol maythen produce diethyl carbonate, which may be transesterified anddisproportionated as above to form diaryl carbonate. Each of thesereactions may be performed in one or more reaction zones, intermediateor inclusive of which may be one or more separation stages forseparating reaction products, reactants, and/or reaction by-products.

Referring now to FIG. 19, a simplified block flow diagram of anintegrated process for the production of diaryl carbonates according toembodiments disclosed herein is illustrated. As illustrated in FIG. 19,a diaryl carbonate, such as DPC, is produced from CO₂, an epoxide, andphenol.

Carbon dioxide and an epoxide, such as ethylene oxide or propyleneoxide, are introduced to reaction zone 203 via flow lines 201 and 202 toproduce a cyclic carbonate such as ethylene carbonate or propylenecarbonate. Cyclic carbonate stream 204 from the synthesis zone 203 andethanol via stream 207 are introduced to a catalytic distillationreactor system 205 to perform transesterification with ethanol in thepresence of a transesterification catalyst. Fresh make-up ethanol stream206, if necessary, may be combined with ethanol recycle streams 220and/or 210 for feed to the reaction zone 205 via stream 207. An overheadstream 208 from reaction zone 205 is sent to a first separation zone209, which may be a distillation column, for example, to separate theproduct DEC from ethanol. The overhead ethanol stream 210 from theseparation zone 209 is recycled back to the transesterification zone205. The DEC stream 211 (bottom liquid stream) from first separationzone 209 is fed to a second transesterification reaction zone 217.

The bottoms flow stream 212 from the transesterification zone 205 issent to the second separation zone 213, which may comprise one or moredistillation columns, to separate the product glycol from unconvertedcyclic carbonate and reaction intermediates. The product glycol from theseparation zone 213 is removed via line 214. The remaining heavierliquid stream 215, including cyclic carbonate, is recycled back to thefirst transesterification zone 205.

Fresh phenol stream 216 is introduced to the second transesterificationzone 217, which comprises at least one catalytic distillation reactorsystem, to concurrently perform the transesterification of DEC withphenol in the presence of a catalyst to produce EPC and ethanolco-product and to separate the EPC from ethanol and unreacted DEC. Theoverhead stream 218, which contains ethanol, DEC, lights and a smallamount of phenol, from the transesterification zone 217 is sent to thethird separation zone 219. The separation of ethanol and DEC containedin stream 218 does not require extractive distillation, as mentionedabove.

Ethanol stream 220 is recycled to the first transesterification zone 205via line 207. Lights are vented through line 221. The bottom stream 222,which contains DEC and phenol, is recycled back to the secondtransesterification zone 217 via line 222. The bottom stream 223 fromthe second transesterification zone 217, which contains the product EPC,phenol and small amounts of DPC, is sent to the EPC disproportionationzone 224, which comprises at least one catalytic distillation reactorsystem operated under a partial vacuum. The overhead stream 225 from thereaction zone 224, which contains DEC, phenol and a small amount ofethanol, is recycled back to the top of the catalytic distillationcolumn of the transesterification zone 217. The bottom stream 226 fromthe reaction zone 224 is sent to the fourth separation zone 227, whichis comprised of two or more distillation columns for the separation ofthe materials in the stream 226. The stream 228 from the separation zone227, which contains unconverted EPC and a trace amount of phenol isrecycled back to the top of the catalytic distillation column of thedisproportionation reaction zone 224.

The final product DPC stream 229 from the separation zone 227 is sent toa DPC storage tank. Heavies in stream 226 are removed from the zone 227via line 230 for disposal or further treatment, if necessary. As afurther option, glycol recovered from separation zone 213 may bedehydrated to form an epoxide and water in a dehydration reactor 232,where the epoxide may then be recycled to form a cyclic carbonate, asdescribed above. The result of recycling glycol is an essentiallyco-product-free DPC process. Under certain circumstances, co-productfree processes are desirable, such as where an outlet for the glycol isnot available.

Referring now to FIG. 20, a simplified block flow diagram of anintegrated process for the production of diaryl carbonates according toembodiments disclosed herein is illustrated, where like numeralsrepresent like parts. A diaryl carbonate, such as DPC, is produced fromCO₂ and phenol according to an alternative method. Ammonia is involvedin this alternative method to produce the intermediate vehicle DEC,which differs from the process route illustrated in FIG. 19.

Carbon dioxide stream 240 and ammonia stream 260, which is a combinedstream of an ammonia make-up stream 241 and ammonia recycle stream 250,are introduced to urea synthesis unit 242. The co-product H₂O of theurea synthesis is removed via line 243 from the synthesis unit 242. Theproduct urea stream 244 from the synthesis zone 242 and ethanol feedstream 247, which is a combined stream of ethanol recycle stream 253 andan ethanol make-up stream 246, are introduced to the urea alcoholysisreaction zone 245.

The urea alcoholysis reaction in zone 245 is a two-step reaction. In thefirst step, the first alcoholysis of urea is performed with ethanol toproduce ethyl carbamate (C₂H₅O—CONH₂) and to remove impurities such asammonium carbamate with co-product ammonia, usually in the absence of acatalyst. The urea may be dissolved in ethanol and the resulting ureasolution is pumped into a reactive distillation column. At the sametime, a super heated ethanol vapor is introduced into the bottom sectionof the column to strip off the ammonia co-product from the liquid phaseinto vapor phase to remove ammonia as a part of overhead stream. Thebottoms fraction recovered from the column comprises ethyl carbamate, aminor amount of urea, and ethanol. In the second step, the secondalcoholysis of ethyl carbamate and remaining urea is performed withethanol in the presence of a catalyst to produce DEC and co-productammonia in another reactive distillation column.

The product streams 248 from the urea alcoholysis zone 245 are sent tothe separation zone 249, which is comprised of one or multipledistillation columns. The ammonia from the separation zone 249 isrecycled via lines 250 and 260 to the urea synthesis zone 242. Therecovered ethanol stream 251 is recycled back to the urea alcoholysiszone 245 via line 251. DEC from the separation zone 249 becomes the DECfeed stream 252 to perform the transesterification with phenol in thereaction zone 217. The co-product ethanol of DEC transesterificationwith phenol in the reaction zone 217 is recycled back to the ureaalcoholysis zone 245 to produce DEC via lines 253 and 247.

The remaining process steps to produce DPC from this point are identicalto the description given above for FIG. 19. Alternatively, urea may bepurchased from a urea producer and return of ammonia co-product for acredit. This may be cheaper in terms of urea cost and more effective interms of energy consumption than operating a small on-site ureasynthesis unit.

As mentioned above, the processes for production of DPC as describedwith respect to FIGS. 19 and 20 may be run with essentially no freshethanol feed. The ethanol is initially consumed during thetransesterification of urea or cyclic carbonates to produce DEC, andthen ethanol is produced during the transesterification of DEC with anaryl hydroxyl compound, such as phenol, to produce a diaryl carbonate,such as DPC. Because the ethanol is consumed and produced atapproximately the same mole ratio, excluding consumption of ethanol inreaction by-products, the processes disclosed herein may be operatedusing an essentially closed loop process with regard to ethanol. Assuch, ethanol raw material and pre-conditioning (drying) costs may besubstantially reduced as compared to a stand-alone DEC process.

Reaction zones for conducting the above transesterification anddisproportionation reactions may contain one or multiple solidcatalysts, which may be contained in one or more reactors. The reactorcan have any physical shape for various operational modes to perform thetransesterification in liquid phase or in the presence of a dual phaseof liquid and vapor. Any type of reactors may be used to carry out thereactions described herein. The examples of reactors suitable forcarrying out the reactions involving organic carbonate or organiccarbamates reactions may include distillation column reactors, dividedwall distillation column reactors, traditional tubular fixed bedreactors, bubble column reactors, slurry reactors equipped with orwithout a distillation column, pulsed flow reactors, catalyticdistillation columns wherein slurry solid catalysts flow down thecolumn, or any combination of these reactors.

Catalyst useful for the transesterification and disproportionationreactions for producing DEC, EPC and DPC in these integrated processesmay be as described above, including one or more of homogeneouscatalyst, heterogeneous catalyst, and solid catalyst according toembodiments disclosed herein and detailed above. Heterogeneous catalystcan have any physical form, such as, for example, a powder for stirredtank reactor operation or a slurry catalytic distillation columnreactor, or a shaped material such as spheres, granules, pellets,extrudates, woven cloth, mesh, etc.

EXAMPLES

All experimental reactions were performed in a catalytic distillationreaction system containing a single-stage vertically-mounteddistillation column having dimensions of 1.3 cm (½ inch) diameter and6.5 cm (25 inches) length, a fixed catalyst bed, and operating as anup-flow boiling point reactor, where vapor and liquid phases coexist.The reactor had separately controlled top and bottom heating zones. Thevolume of solid catalysts was 25 ml.

The following experimental examples illustrate embodiments of producingdiphenyl carbonate (DPC) from either diethyl carbonate (DEC) and phenolor from carbon dioxide and phenol, without the need for a solventextractive distillation of an azeotrope, according to the presentdisclosure.

Example 13

This example illustrates an embodiment for producing diethyl carbonate(DEC) according to the present disclosure as shown in FIG. 19.Transesterification of propylene carbonate (a cyclic carbonate) withethanol in the presence of a solid catalyst to produce DEC was performedat various conditions listed in Table 8 below.

TABLE 8 Ethanol-to- propylene Time on carbonate Stream in feed Feed RateMg in feed, (hrs) (mole ratio) (ml/min) (wtppm) Notes  0-79 6.96 0.5 1111^(st) Transesterification  79-175 10.12 0.5 111 1^(st)Transesterification 175-341 6.95 0.5 56 1^(st) Transesterification341-413 6.89 0.5 34.8 1^(st) Transesterification 413-485 7.17 0.5 42.81^(st) Transesterification 485-570 3.41 0.5 41.5 1^(st)Transesterification 570-666 4.14 0.5 34.5 1^(st) Transesterification666-738 5.39 0.5 34.5 1^(st) Transesterification 738-834 6.59 0.5 37.52^(nd) Transesterification

The starting solid transesterification catalyst was MgO supported onsilica gel. A magnesium nitrate solution was prepared by dissolving 10.1g of Mg(NO₃)₂.6H₂O in 22.7 g of deionized water. Magnesium nitrate wasdeposited on 30 ml (11.8 g) of silica gel support by incipientimpregnation. The silica gel support (˜3 mm diameter spheres) had 314m²/g BET surface area, 1.06 cm³/g pore volume, and 13.46 nm average porediameter. The impregnation product was dried at 100° C. in vacuum ovenfor one hour, followed by calcinations at 510° C. for 2 hours, toprepare MgO supported on silica gel. 25 ml (10.77 g) of MgO catalystsupported on silica gel was loaded in the reactor.

Magnesium tert-butoxide was also added to the reaction feed solutions invarious trace amounts in order to obtain stable catalyst activity, asshown in Table 8 above. In general, addition of a trace amount of asoluble organo alkaline earth compound, such as an alkoxide, aglycoloxide, or a mixture thereof, can be used to obtain extendedcatalyst cycle time and stability, as described above for embodimentsdisclosed herein.

The transesterification reaction products included 1-ethoxy-2-propanoland dipropylene glycol as by-product. No diethyl ether was detected inany product sample. Results for Experiment 13 are also illustrated inFIG. 21. The average selectivities for forming DEC and propylene glycolin the first transesterification were 95.3 mole percent and 94.8 molepercent, respectively. The selectivities declined slowly with conversionof propylene carbonate and increased with an increase in the mole ratioof ethanol to propylene carbonate. For example, the averageselectivities of DEC and propylene glycol for the secondtransesterification were 94.0 mole percent and 92.8 mole percent,respectively.

Example 14

This example illustrates an embodiment for producing diethyl carbonate(DEC) according to the present disclosure as shown in FIG. 20.Alcoholysis of ethyl carbamate with ethanol to produce DEC and ammoniaco-product in the presence of a solid catalyst was performed at variousconditions listed in Table 9 below:

TABLE 9 Dibutyl Tin Time on Stream Temperature Pressure Feed Rate N2Flow Rate Dimethoxide (hrs) (° C. (° F.)) (bar (psig)) (ml/min) (cc/min)(wtppm Sn)  0-72 174 (345) 6.2 (75) 0.5 1.5 650 72-96 174 (345) 6.2 (75)0.5 1.5 1300  96-156 174 (345) 6.2 (75) 0.5 1.8 1300 156-254 174 (345)6.2 (75) 0.5 1.5 1950 254-298 174 (345) 5.4 (63) 0.5 0 1950 298-406 174(345) 5.4 (63) 0.5 1.0 1950 406-456 174 (345) 5.4 (63) 0.5 1.5 1950

The solid catalyst was prepared by immobilizing dibutylin dimethoxide ona silica gel by in situ technique. To prepare the catalyst, 25 ml (14.79g) of spherically shaped silica gel support having an approximatediameter of 3 mm were loaded in the reactor. The silica gel support hadabout 6 hydroxyl groups per nm², a BET of 392 m²/g, a pore volume of0.633 cm³/g, an average pore diameter of 6.48 nm, and an ABD of about0.58 g/ml.

A dibutyl dimethoxide solution was prepared by mixing 87 g of dibutylindimethoxide in 2 liters of dry toluene. The reactor was filled with thissolution up-flow at ambient temperature and pressure. The reactor washeated slowly to a temperature of 110° C. (230° F.) while adding thedibutyl dimethoxide solution up-flow at a rate of 2 ml/min. At 110° C.(230° F.), the reactor was placed under a pressure of 3.4 bar (35 psig)and then was heated to 135° C. (275° F.).

While maintaining the reactor temperature at 135° C. (275° F.) andreactor pressure at 3.4 bar (35 psig), the dibutylin dimethoxidesolution was passed through the reactor up-flow at 0.5 ml/min for 6hours. After cooling, excess solution in the reactor was drained andthen the catalyst was washed with dry toluene at 4 ml/min up-flow for1.5 hours. The washed catalyst was dried at a temperature of 104° C.(220° F.) and ambient pressure under a nitrogen down-flow purge of 300cc/min for a duration of 2 hours.

The alcoholysis reaction of ethyl carbamate with ethanol to produce DECwas performed by passing a solution of 13.2 weigh percent ethylcarbamate, 31.36 weight percent triglyme and 55.44 weight percentethanol up-flow with nitrogen gas through the solid catalyst bed asdescribed above under boiling point conditions. One may also carry outthe reaction in down-flow reactor. Trace amounts of dibutylindimethoxide were blended into this solution during the reaction.

The reaction conditions are listed in Table 9 above, and the results ofthis experiment are illustrated in FIG. 22. The selectivity of DEC basedon ethyl carbamate was in a range of from 98.5 mole percent to 99.9 molepercent with a general trend of decreasing selectivity as the conversionof ethyl carbamate increased.

Example 15

This example illustrates an embodiment for producing ethyl phenylcarbonate (EPC) according to the present disclosure as shown in FIG. 19.Transesterification of DEC with phenol to produce ethyl phenyl carbonate(EPC) was performed in the presence of titanium ethoxide catalystimmobilized on a silica gel support.

The same silica gel support as in Example 14 above in the amount of 25ml (14.47 g) was loaded in the reactor, followed by immobilization oftitanium ethoxide. A titanium ethoxide solution was prepared bydissolving 45.25 g of titanium ethoxide in 800 ml of toluene. Thetitanium ethoxide solution was then circulated up-flow through thereactor at a rate of 15 ml/min and at ambient temperature and pressurefor a duration of 30 minutes. The titanium ethoxide was then graftedonto the silica gel support at 135° C. (275° F.) and 3.4 bar (35 psig)for 17 hours. After cooling, the excess titanium ethoxide solution wasdrained out of the reactor and the catalyst was washed with toluene at arate of 4 ml/min for 1.5 hours. The washed catalyst was dried at 138° C.(280° F.) for 4 hours under a nitrogen gas flow of 300 cc/min.

The transesterification of DEC with phenol was performed to produceethyl phenyl carbonate (EPC) under the boiling point reactor conditionof 345° C. and 27 psig by pumping 0.5 ml/min of 23 weight percentsolution of DEC into the reactor.

To stabilize catalyst performance during the entire 860 hours ofon-stream time, titanium in the form of Ti (Oet)_(4−x)OPh)_(x) (wherex=2) was introduced in the feed stream at a concentration of about 55wtppm. The combined selectivity of the transesterification reactiontowards EPC and DPC was about 99 mole percent. The result of thisexperiment is illustrated in FIG. 23.

Example 16

This example illustrates an embodiment for producing diphenyl carbonate(DPC) according to the present disclosure as shown in FIG. 19.Disproportionation of EPC to produce DPC and DEC was performed in thepresence of soluble titanium catalyst.

The feeds for the disproportionation reaction were prepared bydistilling ethanol, DEC and a part of phenol from a compositetransesterification product under a nitrogen blanket, which was producedby the experiments similar to the one disclosed in Example 15 above. Thehomogeneous titanium catalyst in the feed mixture to thedisproportionation reaction was as originally introduced in thetransesterification reactor, as disclosed in Example 15 above. Noadditional soluble titanium catalyst was added to the feed mixtures.Toluene was added to two feed mixtures to create a vapor phase for theboiling point reactor.

The first feed composition was 16.26 weight percent toluene, 1.61 weightpercent DEC, 49.33 weight percent phenol, 30.91 weight percent EPC, 0.78weight percent DPC, and the net of the composition balance accountingfor by-products from the transalkylation reaction such as MPC. Thesecond feed composition was 16.15 weight percent toluene, 1.61 weightpercent DEC, 49.28 weight percent phenol, 31.08 weight percent EPC, 0.80weight percent DPC, and the net of the composition balance accountingfor by-products from the transalkylation reaction such as MPC. Theconcentration of homogeneous catalyst in the first and second feeds was180 weight ppm and 200 weight ppm titanium, respectively.

The disproportionation reaction was performed in the reactor with 25 mlof empty catalyst space (in the absence of solid catalyst) at 179° C.(355° F.) and 2.9 bar (27 psig). The feed rates were 0.5 ml/min up-flowfor the first 72 hours on stream and then 0.60 ml/min up-flowthereafter. The results of the disproportionation reactions areillustrated in FIG. 24. Xanthone was the only new by-product producedduring disproportionation in an amount of about 35 wtppm. Diphenyl etherwas not detected in any product sample analyses. The selectivity of allthe by-products was from 3.0 mole percent to 3.3 mole percent.

Production of Biodiesel According to Embodiments Disclosed Herein

Biodiesel has been produced by performing transesterification ofvegetable oils and animal fats with methanol in the presence ofhomogeneous catalysts and solid catalyst. Feedstocks for the productionof bio-diesel are vegetable oils and animal fats, which are esters ofhigher fatty acids. The term fat (vegetable or animal oil, if liquid) isusually confined to esters (glycerides) of fatty acids with glycerol,and the term wax to esters of other alcohols. The basic chemistryinvolved in producing bio-diesel is the catalytic exchange reaction ofnatural esters (mainly glycerides) with a primary alcohol (typicallymethanol or ethanol). An alcoholic solution of a base (usually NaOH,KOH, potassium methoxide, or sodium methoxide) may be used as acatalyst. Therefore, bio-diesel is a mixture of methyl or ethyl estersof various saturated and unsaturated fatty acids. Co-product isglycerol, which amounts from 16 to 25 wt %. Biodiesel may also containsome fatty acids (hydrolysis products of esters) in minor amountsdepending on the amount of water in the feed or the catalyst used.

-   -   where R⁴OH=methanol or ethanol; R=R¹, R² or R³.

The alkyl groups R¹, R² and R³ of the natural product glyceride are, ingeneral, different in the chain length and degree of unsaturation. Thealkyl groups are usually straight chain and have even number of carbonatoms from 4 to 26. The exception is branched isovaleric acid(CH₃)₂CHCH₂COOH, which occurs in relatively large amounts in dolphins.Some unsaturated fatty acids have two or three double bonds in the alkylchains. Unsaturated fatty acids have lower melting points than theirsaturated counter parts. The chain length of unsaturated fatty acids isgenerally in the range of C₁₀-C₂₄. Canola oil has a higher degree ofunsaturation in C₁₆-C₂₀ chain length than corn oil.

In general, base catalysts are more effective for transesterification ofcarboxylic esters with an alcohol than acid catalysts. The heterogeneouscatalysts disclosed in the prior art (see Background) are also basecatalysts. Unfortunately, active catalytic components leach out of solidcatalyst under the reaction condition, resulting catalyst deactivation.Zinc aluminate catalysts are not very active catalyst and require higherreaction temperatures and lower feed rates than more basic catalystssuch as MgO or CaO. But the latter leach out of the solid catalyst evenfaster than the zinc aluminates.

Transesterification of vegetable oils or animal fats may be carried outwith methanol or ethanol in the presence of a solid catalyst in aboiling point reactor, pulse flow reactor, or catalytic distillationcolumn with a trace amount of soluble catalytic component in the feedmixture in a one step or two step reaction. Starting catalysts mayinclude metal oxides, such as magnesium oxide, calcium oxide, zincoxide, sodium oxide, potassium oxide, lanthanum oxide, etc., supportedon a support, such as silica, alumina carbon and/or carbonaceousmaterial. Carbon and carbonaceous supports will preferably have surfacefunctional groups such as hydroxyl or carbonyl or both to immobilizeorganometallic compounds on the surface of the support.

To prepare supported metal oxides, hydroxides, or oxyhydroxides, thesurface functional groups may not be necessary. Carbonaceous supportsmay be prepared by controlled thermal dehydration of carbohydrates, suchas wood, coconut shell, starch, cellulose, a mixture of starch andcellulose, sugar, methyl cellulose, etc., at elevated temperatures.Carbonaceous supports may be either unsupported or supported. To preparesupported carbonaceous material, carbohydrates are deposed on a suitableporous support followed by controlled thermal dehydration at an elevatedtemperature from 300° C. to 1000° C. in an inert atmosphere or anatmosphere composed of an inert gas, a small amount of oxygen or steamor both. Support for carbonaceous materials may be any inorganicmaterials such as alumina, titania, silica, zirconia, clays,silica-alumina, etc.

In the two step process, the conversion of triglyceride across the firstreactor may be higher than about 90%. Remaining unconvertedtriglyceride, diglyceride and monoglyceride in the reaction productstream from the first transesterification reactor may be converted tocompletion in the second transesterification reactor. Since thetransesterification is a two phase reaction, performingtransesterification in a boiling point or pulsed flow reactor will aidto transport large triglyceride molecules, methyl esters, and stickyglycerol through catalyst pores back and forth between bulk liquidmedium and interior of catalyst pellets, where most catalytic reactionsoccur, resulting in high productivity. Since catalysts disclosed hereinhave high activity, the transesterification may be performed at a lowertemperature and pressure, which means lower construction cost andutility cost.

The addition of soluble catalytic component to the feed stream to thereactors is from about 0.5 ppm to about 500 ppm by weight in someembodiments; from about 5 ppm to about 250 ppm by weight in otherembodiments; and from 10 ppm to 50 ppm by weight in other embodiments.Examples of soluble catalytic compounds include zinc 2-methoxyethoxide,calcium 2-methoxyethoxide, zinc 2-methoxypropoxide, zinc ethoxide, zincalkoxy alkyl carbonate, calcium 2-methoxyproxide, calcium ethoxide,calcium methoxide, calcium alkoxy alkyl carbonate, magnesium2-methoxyethoxide, magnesium 2-methoxyproxide, magnesium ethoxide,magnesium methoxide, magnesium butoxide, magnesium alkoxy alkylcarbonate, lanthanum alkoxide, lanthanum alkoxy alkyl carbonate, zincsalts of carboxylic acids, magnesium salts of carboxylic acids, calciumsalts of carboxylic acids, and Mg, Ca, and Zn glycerides, among others.A mixture of these may also be used. Soluble compounds of Ca, Mg, Zn andLa may be obtained by reacting oxide or hydroxide of these metal with anorganic carbonate or a mixture of organic carbonate and an alcohol, orcarboxylic acids or a mixture of organic carboxylic acid and an alcoholsuch as methanol, 2-methoxyethanol, etc. at temperature from 93° C. to260° C. (200° F. to 500° F.), preferably from 121° C. to 232° C. (250°F. to 450° F.) in liquid phase or presence of liquid and vapor.Optionally one may choose to recover the metal components for recycle.Such prepared solutions are useful for adding trace amount of thesemetal into the feed stream to a reactor to obtain a long catalyst cycletime. Total amount of active metal or metal components on a solid metalalkoxide, metal hydroxide or metal oxide catalyst is from about 0.05 wt% to about 20 wt % in some embodiments, and from about 0.07 wt % toabout 12 wt % in other embodiments.

Optionally, all or a part of di- or mono-glycerides may be converted toorganic carbonates or organic carbamates or both by reaction with DMC,methyl 2-ethyl-1-hexyl carbonate, methyl carbamate, 2-ethyl-1-hexylcarbamate, urea, or a mixture of these in addition totransesterification with methanol in the second reactor or optionally ina third reactor. The resulting organic carbonates and carbamates mayserve as a biodiesel additive agent for reducing particulates, NO_(x)emissions, or improvement of diesel cetane.

As natural vegetable oils may contain various minor amounts of freefatty acids, free fatty acids needed to be removed by pretreatment priorto performing transesterification with an alcohol in the presence of asolid base catalyst. An example of such pretreatment methods isesterification of free fatty acids with methanol in the presence of anacid catalyst. One such acid catalyst is sulfonic acid immobilized on acarbonaceous support. Support may include those prepared by controlledthermal dehydration of coconut shell or carbohydrates supported ordeposited on a porous support. Performing esterification of free fattyacids with an alcohol in the presence of a solid acid catalyst in acatalytic distillation reactor has advantages, which are continuousremoval of water from the reaction zone as an overheads stream, drivingthe esterification toward completion and eliminating a separate dryingstep of the esterification product prior to performingtransesterification of triglyceride with an alcohol. Another importantadvantage is decreased esterification times.

All transesterification reactions in the following examples wereperformed in down-flow reactors. The dimension of the fixed bed reactorwas 1.3 cm (½ inch) diameter by 53.3 cm (21 inches) long. The reactorhad separately controlled top and bottom heating zones. The feedmethanol stream and vegetable oil stream (6 wt % methanol in vegetableoil) were separately pumped into the top section of the reactor, wherethe two streams flow down into the catalytic reaction zone. Traceamounts of soluble catalytic component were blended into a methanolstream or already contained partially converted product stream. Thevolume of solid catalysts was 15 ml.

Experiment 17

The objective of this experiment was to demonstrate thetransesterification of canola oil with methanol in the presence of asolid catalyst in a down-flow boiling reactor or catalytic distillationreactor. The solid catalyst is MgO supported on a silica gel.

A magnesium nitrate solution was prepared by dissolving 10.96 gMg(NO₃)₂.6H₂O in 24 g deionized water. 30 ml (11.91 g) of a silica gelsphere support (1.7-4 mm diameter; about 6 hydroxyl groups per nm², 314m²/g BET, 1.055 cm³/g pore volume and 13.46 nm average pore diameter)was impregnated with the above magnesium nitrate solution by anincipient wetness technique. The silica gel sphere support was preparedby an oil dropping technique. After drying the impregnation product at100° C. for 1 hour, it was then calcined at 510° C. for 2 hrs.

15 ml (6.30 g) of the MgO/SiO₂ catalyst was loaded in the reactor.Canola oil feed (purchased from a local grocery store) was prepared bymixing methanol (5.39 wt %) with the canola oil (94.61 wt %). The acidvalue of free fatty acid of this feed was 0.48 mg KOH/g. Thetransesterification of canola oil with methanol was performed at 165° C.(330° F.) and 19.4 bar (267 psig) by feeding the canola oil feed andmethanol at 0.2 ml/min each. Magnesium ethoxide was dissolved in themethanol feed to have 28 ppm Mg by weight in the catalytic reactionzone.

The effluent streams were composed of two clear layers. Top layerscontain the product methyl esters and small amounts of unconvertedtriglycerides. The average content of unconverted triglycerides in thereaction products, excluding methanol from the top layers, was about 1.2wt %. The bottom layer contains most of the unconverted triglycerides.The result is illustrated in FIG. 17, which indicates a stable catalystperformance.

Experiment 18

The objective of this experiment was to demonstrate converting theremaining unconverted or partially converted materials in the effluentstream (two layers on standing) from the first reactor, in a seconddown-flow boiling point reactor or catalytic distillation reactor, oroptionally recycle to the front of a single transesterification reactor.

A magnesium nitrate solution was prepared by dissolving 9.04 gMg(NO₃)₂.6H₂O in 19.1 g deionized water. 22 ml (9.14 g) of a silica gelsphere support (9-14 mesh, 309 m²/g BET and 1.03 cm³/g pore volume) wasimpregnated with above magnesium nitrate solution by incipient wetnesstechnique. After drying the impregnation product at 150° C. for 1 hour,it was calcined at 510° C. for 2 hrs. The fished catalyst contained 4.5%Mg by weight.

15 ml (7.2 g) of the MgO/SiO₂ catalyst was loaded in the same reactorused in Experiment 13. The two layers of a composite product from thefirst transesterfication reaction of canola oil with methanol wereseparated from the composite product by using a separation funnel to beused as feeds for the second transesterfication reaction. Thecomposition of the bottom composite product feed was 25.4 wt %triglycerides, 8.5 wt % diglycerides, 3.1 wt % monoglycerides, 0.1 wt %glycerin, 47.1 wt % methyl esters and 15.8 wt % methanol. The feedcontained about 8.5 ppm soluble Mg species by weight and had 0.32 mgKOH/g free fatty acid value. Transesterification was performed at 160°C. (320° F.) and 19.5 bar (268 psig) by pumping 0.12 ml/min feed and0.10 ml/min methanol into the down-flow boiling point reactor. Noadditional Mg alkoxide was added into either of the two feed streams.The reactor effluent stream was a clear light yellow solution (singlelayer).

The composition of the top composite product feed was 1.12 wt %triglycerides, 0.57 wt % diglycerides, 3.78 wt % monoglycerides, 7.47 wt% methyl esters, 0.03 wt % glycerin and 87.03 wt % methanol. The freefatty acid value of this feed was 0.51 mg KOH/g. The transesterificationwas performed over the same catalyst at the same temperature andpressure at a 0.2 ml/min feed flow rate. No additional methanol waspumped into the reactor. These two final transesterification productsfrom composite bottom and top composite product feeds were combined todistill off excess methanol and to recover crude biodiesel. Therecovered crude biodiesel contained 0.36 wt % unconverted triglyceridesand 0.74 mg KOH/g free fatty acid value.

The above experimental result successfully demonstrates that biodieselcan be produced by performing transesterification of vegetable oil withan alcohol such as methanol in the presence of a solid catalyst

Catalyst Systems

As described for various embodiments in the present specification,reaction zones or catalyst beds in reactors may include one or morealcoholysis catalysts according to embodiments disclosed herein,including reaction zones or catalyst beds having two or moreheterogeneous catalysts, solid catalysts, or combinations thereof. Forexample, a catalyst may include two or more metal oxides, i.e., a mixedmetal oxide catalyst having two or more different metal oxides depositedor immobilized on a support, which may also be a metal oxide or a mixedmetal oxide that is porous.

Solid or heterogeneous catalysts as used herein refers not just to theactive metal components of a catalyst, but also to the properties of thesupport, which by themselves may distinguish a catalyst. For example,two different catalyst supports, which have different physicalproperties such as crystallinity, porosimetry, density, size, etc., buthave the same chemical composition, are defined herein as two differentsupports.

Similarly, solid or heterogeneous catalysts may vary based on theconcentration of the active component on a support, such as 1 wt. %metal oxide versus 3 wt. % metal oxide on a support. While each may havea similar metal oxide and a similar support composition and structure,the metal loading may result in substantially differing catalystbehavior.

Thus, the “two or more” catalysts used in embodiments disclosed hereinmay include catalysts having a multiple active catalytic components,different support structures/compositions, different loadings of activecomponents, or any combination or permutation of metal elements, metalloadings, and supports.

Reaction zones having two or more catalysts according to embodimentsdisclosed herein may advantageously exploit the differingcharacteristics of the catalysts. For example, a catalytic distillationreaction zone may be tailored to contain multiple catalysts, providingfor the desired conversion with minimal polymerization or fouling. Forexample, catalysts having a lower activity and/or high selectivity maybe used in portions of the bed where the concentration of polymerprecursors may be the highest. Alternatively, catalysts having a largeaverage pore diameter or other properties advantageously allowing theliquid traffic in the distillation column reactor to wash the impuritiesfrom the catalyst may be used toward the bottom of a reaction zone wherethe concentration of polymer precursors may be the highest. Catalystshaving a high activity or high surface area may be selectively locatedat portions of the bed where polymer precursor concentration is low. Inthis manner, the catalyst bed may be tailored to provide for the desiredconversion (activity and selectivity) and low or no fouling of thecatalysts within the bed, extending catalyst cycle time and enhancingreactor performance.

Water

As mentioned above, the moisture content of the feed stream may becontrolled to be less than about 700 ppm in some embodiments, and lessthan about 600 ppm in other embodiments. It has unexpectedly been foundthat use of trace amounts of water in the feed stream or in the reactionzone may improve catalyst cycle time and/or reduce build-up of heavypolymeric materials on the solid catalyst, which may result in slowerdeactivation rates.

Not being bound to a single theory, it is noted that the mechanism ofimproved catalyst performance using heterogeneous and homogeneouscatalysts in the presence of water is not fully understood. It is notedhowever, that (a) deposition of polymers on the catalyst deactivates thecatalyst, and (b) leaching of active metals our of the solid catalystunder reaction conditions causes permanent catalyst deactivation. It istheorized that the use of trace amounts of water may (a) result indepolymerization of polymers that may form during thetransesterification and/or disproportionation reactions in embodimentsdisclosed herein, and/or (b) participate in the reaction to anchor,deposit, or tether one or more of leached catalysts or added homogeneouscatalysts to the support at reaction conditions. A trace amount of waterin an amount of less than about 600 ppm by weight in the feed or addedby directly or indirectly injecting water into the reaction zone mayresult in the effect of depolymerization, in-situ catalyst reactivation,or a combination thereof.

Too much water in feed stream or reaction zone, however, must beavoided, as this may cause precipitation of soluble catalyst orformation of gel on solid catalyst or both, which may cause undesirableproblems associated with reactor operation and optimal catalystperformance. The trace of amount of water in feed stream or a reactionzone in embodiments disclosed herein may be in a range from about 1 ppmto about 600 ppm; from about 2 ppm to 500 ppm in other embodiments; fromabout 5 ppm to about 400 ppm in other embodiments; and from about 50 ppmto about 250 ppm in yet other embodiments, where each is by weight.

Referring again to FIGS. 19 and 20, water may be produced as a reactionby-product in each flow scheme for the production of diaryl carbonatesaccording to embodiments disclosed herein. With regard to the process ofFIG. 19, water may be produced during the dehydration of glycol to forman epoxide. With regard to the process of FIG. 20, water may be producedduring the synthesis of urea from ammonia and carbon dioxide. Waterrecovered from either of these reaction steps may be used as added waterto one or more transesterification and/or disproportionation reactionzones to maintain the water within the respective reaction zones in theabove described ranges.

Internal Devices Used in Reactors

As described above, reactors used in embodiments disclosed herein canhave any physical shape for various operational modes to perform thetransesterification in liquid phase or in the presence of a dual phaseof liquid and vapor. Any type of reactors may be used to carry out thereactions described herein. The examples of reactors suitable forcarrying out the reactions involving organic carbonate or organiccarbamates reactions may include distillation column reactors, dividedwall distillation column reactors, traditional tubular fixed bedreactors, bubble column reactors, slurry reactors equipped with orwithout a distillation column, pulsed flow reactors, catalyticdistillation columns wherein slurry solid catalysts flow down thecolumn, or any combination of these reactors.

Reactors used in embodiments disclosed herein may additionally includeany internal physical device or a combination of two or more internaldevices for vapor-liquid separation and directing of vapor-liquidtraffic within the reactor. The various internal devices can be anyshape or physical device and may have multiple purposes as long as suchdevice or devices promote both effective vapor-liquid separation andvapor-liquid traffic.

In dealing with an endothermic equilibrium reaction, achievingsimultaneously both tasks of effective vapor-liquid separation andvapor-liquid traffic is necessary for high conversion and yet iscomplicated by the cooling effect of the reaction medium due to twocombined effects of heat of vaporization and heat of reaction. Toachieve high conversion, one of the reaction products has to be removedfrom liquid reaction medium to vapor phase and then vapor has to beremoved quickly from the reaction zone, which require heat ofvaporization and effective vapor traffic. An example of such a case istransesterification of a dialkyl carbonate such as DEC with phenol. Theresult is cooling of the liquid reaction medium, which in turn resultsin lower conversion. Heat of vaporization and endothermisity of reactionare conflicting, resulting in lower conversion. In traditional tubularplug flow reactor, one may use internal heating system or use multiplesmall reactors with intermittent heating between reactors.

One may be able to use an internal auxiliary device or devices in thereaction zone, which effectively separates gas and liquid, and alsosupply heat for the endothermic reaction. Such device may be liquidredistribution tray with a device or devices for catalytic distillationcolumn, which supplies heat internally to the middle of reaction zone.Such a device or devices should promote more effective separation ofvapor from liquid as well as supply heat for an endothermic reaction,resulting in high conversion. Transesterification of a dialkyl carbonatesuch as DEC with phenol or a cyclic carbonate with ethanol shouldbenefit from use of such device or devices in the reaction zone ofcatalytic distillation column regardless of use of solid catalyst,homogeneous catalyst, or both.

For exothermic reactions, it may be possible to use an internal coolingdevice to prevent a run-away reaction or drying of the reaction zone(low liquid traffic), which usually results in poor selectivity and/orpoor catalyst performance. The internal device has to be beneficial forseparation of vapor-liquid phase and promote vapor-liquid traffic.

The internal heating or cooling devices may be placed at any heightwithin the column, such as at a top of a rectification section of acolumn, such as a transesterification reactor, partially condensing DECand mostly condensing phenol within the column.

As described above, embodiments disclosed herein provide for extendedcatalyst cycle times for various solid catalysts through theintroduction of a trace amount of soluble organometallic compounds withthe feed. Other embodiments disclosed herein may include a method forcontinuously producing organic carbonate or organic carbamate at astable rate; techniques for in situ catalyst preparation of immobilizedsolid catalysts, techniques for maintaining stable catalyst activity forlong catalyst cycle times and service times so as to be suitable forcommercial fixed bed reactors; and an in situ method of reactivatingdeactivated solid catalysts.

Advantageously, embodiments disclosed herein may provide fortransesterification catalysts having an extended cycle length, thereforedecreasing operational costs associated with frequent shut downs andcatalyst changes. Additionally, due to the trace amount of solubleorganometallic compound used, removing the homogeneous catalyst fromvarious product streams may be substantially reduced.

While the disclosure includes a limited number of embodiments, thoseskilled in the art, having benefit of this disclosure, will appreciatethat other embodiments may be devised which do not depart from the scopeof the present disclosure. Accordingly, the scope should be limited onlyby the attached claims.

1. A process for production of diaryl carbonate, comprising: reacting an epoxide and carbon dioxide in a first reaction zone to form first reaction product comprising a cyclic carbonate; transesterifying the cyclic carbonate with ethanol in the presence of a first transesterification catalyst in a second reaction zone to form a second reaction product comprising diethyl carbonate and glycol; separating the second reaction product to recover a first diethyl carbonate fraction and a first glycol fraction; transesterifying at least a portion of the first diethyl carbonate fraction with an aryl hydroxy compound in the presence of a second transesterification catalyst in a third reaction zone to form a third reaction product comprising ethyl aryl carbonate and ethanol; separating the third reaction product to recover an ethyl aryl carbonate fraction and a first ethanol fraction; disproportionating at least a portion of the ethyl aryl carbonate fraction in the presence of a disproportionation catalyst in a fourth reaction zone to form a fourth reaction product comprising diaryl carbonate and diethyl carbonate; separating the fourth reaction product to recover a diaryl carbonate fraction and a second diethyl carbonate fraction; recycling at least a portion of the first ethanol fraction to the second reaction zone; and recycling at least a portion of the second diethyl carbonate fraction to the third reaction zone.
 2. The process of claim 1, further comprising dehydrating at least a portion of the first glycol fraction to form an epoxide and water; separating the water from the epoxide; recycling at least a portion of the epoxide to the first reaction zone.
 3. The process of claim 1, wherein the first transesterification catalyst and the second transesterification catalysts each independently comprise at least one of a solid transesterification catalyst, a soluble organometallic compound, and combinations thereof.
 4. The process of claim 3, wherein the first transesterification catalyst comprises a solid transesterification catalyst, the process further comprising feeding a trace amount of soluble organometallic compound to the second reaction zone.
 5. The process of claim 4, wherein the soluble organometallic compound is fed at a rate in the range of 1 ppm to 200 ppm, based on a total weight of reactants.
 6. The process of claim 3, wherein the second transesterification catalyst comprises a solid transesterification catalyst, the process further comprising feeding a trace amount of soluble organometallic compound to the third reaction zone.
 7. The process of claim 6, wherein the soluble organometallic compound is fed at a rate in the range of 1 ppm to 200 ppm, based on a total weight of reactants.
 8. The process of claim 7, further comprising maintaining a concentration of water in the third reaction zone in a range from 1 ppm to 600 ppm by weight.
 9. The process of claim 3, wherein the solid transesterification catalyst and the soluble organometallic compound each independently comprise at least one Group II to Group VI element.
 10. The process of claim 1: wherein the second reaction zone comprises a catalytic distillation reactor system for concurrently transesterifying the cyclic carbonate with ethanol and separating the second reaction product; wherein the third reaction zone comprises a catalytic distillation reactor system for concurrently transesterifying the diethyl carbonate with the aryl hydroxyl compound and separating the third reaction product; wherein the fourth reaction zone comprises a catalytic distillation reactor system for concurrently disproportionating the ethyl aryl carbonate and separating the fourth reaction product; wherein the first diethyl carbonate fraction comprises diethyl carbonate and ethanol, the process further comprising: separating the first diethyl carbonate fraction to recover a third diethyl carbonate fraction and a second ethanol fraction; and recycling the second ethanol fraction to the second reaction zone; wherein the first glycol fraction comprises glycol and cyclic carbonate, the process further comprising: separating the first glycol fraction to recover a second glycol fraction and a cyclic carbonate fraction; and recycling the cyclic carbonate fraction to the second reaction zone; wherein the diaryl carbonate fraction comprises diaryl carbonate, ethyl aryl carbonate, aryl hydroxyl compound, and reaction byproducts, the process further comprising: separating the diaryl carbonate fraction to recover a diaryl carbonate product fraction, a reaction byproduct fraction comprising compounds heavier than the diaryl carbonate, and at least one recycle fraction comprising at least one of the ethyl aryl carbonate and the aryl hydroxyl compound; and recycling the at least one recycle to the fourth reaction zone; and wherein the first ethanol fraction comprises ethanol and diethyl carbonate, the process further comprising: separating the first ethanol fraction to recover a third ethanol fraction and a fourth diethyl carbonate fraction; recycling the fourth diethyl carbonate fraction to the third reaction zone; and feeding at least a portion of the third ethanol fraction as the at least a portion of the first ethanol fraction recycled to the second reaction zone.
 11. The process of claim 10, further comprising dehydrating at least a portion of the second glycol fraction to form an epoxide and water; separating the water from the epoxide; recycling at least a portion of the epoxide to the first reaction zone.
 12. The process of claim 1, wherein at least one of the second, third, and fourth reaction zones comprise a catalytic distillation reactor system, and wherein at least one catalytic distillation reactor system comprises an internal heating exchange device promoting vapor-liquid separation and directing of vapor-liquid traffic within the catalytic distillation reactor system.
 13. The process of claim 1, wherein at least one of the second reaction zone, the third reaction zone, and the fourth reaction zone comprises two or more transesterification catalysts arranged to promote catalyst longevity and limit catalyst fouling.
 14. The process of claim 13, wherein the two or more transesterification catalysts comprises at least one of different active metal components, different support compositions, and different support properties.
 15. The process of claim 1, wherein the process is operated essentially closed loop with respect to ethanol. 